Vertical integration of source water treatment

ABSTRACT

Conventional oil-water separation methods are inefficient since they break down a given “primary” phase into two “secondary” phases, one is richer and the other one is poorer in the “secondary” phase of the “primary” phase. As such, neither oil is recovered as a readily de-watered stream nor is water recovered as a readily de-oiled stream. However, de-watering and de-oiling of “oil-water” streams are synonymous, and therefore they should be simultaneously targeted by an efficient method. There are provided herein systems and methods to effectively treat “oil-water” streams by simultaneously de-watering the oil phase and de-oiling the water phase, de-scaling the de-oiled water phase, and de-salting the de-scaled water phase.

BACKGROUND OF THE INVENTION Seawater and the Like

Seawater is nearly a uniform saline stream in terms of ions contentexcept seasonal temperature fluctuations and algae bloom. However, theions content may increase in situations where very large amounts ofreject streams from de-salting methods are blown back to a sea,especially in a bay or a semi-closed sea. De-salting methods are notZero-Liquid Discharge (ZLD), but rather feed stream splitters whereinthe ions content in the feed stream is, to a varying degree depending onthe de-salting method, reduced in the product stream but elevated in thereject stream. The discharge of enormous volumes of reject streams fromde-salting methods elevates total dissolved solids (TDS), includingscale prone species, around the seawater intake lines, particularly whenthe dispersion of reject streams is not fast enough. This, in turn, hasa detrimental effect on de-salting methods. Yet, reject streams aredepleted of dissolved oxygen, and enriched with residues of oxygenscavengers, potential toxic species (e.g., derivatives of boron,chlorine and bromine) and gypsum; which adversely affect marineenvironment.

Other forms of reject streams may also result from using seawater influe gas de-sulfurization on: a once-through basis relying on thenatural alkalinity of seawater as a scrubbing agent; a once-throughbasis relying on activating magnesium in seawater as a scrubbing agent;or a recycle-bleed basis utilizing seawater as a makeup water. Suchreject streams are also depleted of dissolved oxygen and enriched withgypsum. Table 1 presents samples of seawater and reject streams fromde-salting and partially de-salting methods including reverse osmosis(RO), multi-stage flash desalination (MSF), and nanofiltration (NF) aswell as a flue gas de-sulfurization method. The term “reject stream” mayalso refer to as “concentrate stream”, “reject brine”, “spent seawater”or “spent water”.

The profound reason for generating staggering volumes of reject streamsfrom, for example, de-salting seawater is the aqueous solubility limitsof the three hydrates of calcium sulfate (dihydrate or gypsum,hemihydrate and anhydrite). The saturation envelops of such hydratesdirectly control the recovery ratio of any de-salting method, causeenormous engineering difficulties, and hinder water production at a lowcost. Within certain temperature confines, such hydrates: (1) are allmetastable; (2) have inverse solubilities with increasing temperaturesexcept gypsum below 40° C.; (3) have solubility patterns in which thesolubility increases with increasing sodium chloride concentrations,reaches a maximum, and then declines at high concentrations of sodiumchloride; yet retaining the same solubility patterns in the presence ofcalcium chloride but decrease with increasing calcium chlorideconcentrations; and (4) require a sufficient detention time (may extendto hours) to induce precipitation as long as water is flowing even whentheir ions pairing conditions and concentrations may seem conducive toprecipitation. Rather than eliminating calcium sulfate hydrates toobviate their profound limitations on any de-salting method, the natureof calcium sulfate hydrates is instead manipulated, to operateinefficient and costly de-salting methods.

For example, pressure-driven membranes (RO and NF) are commonly used totreat seawater and the like at ambient temperatures, and thus thecalcium sulfate would be in the form of gypsum. In such membranes, thereduction-elevation level of ions between a product stream and a rejectstream is governed by the membrane's rejection rate for each ion and thepermissible recovery ratio (product flow rate/feed flow rate). However,the recovery ratio is controlled by the concentration polarization ofrejected ions at the membrane surface as depicted in FIG. 1 [U.S. Pat.No. 7,093,663].

The concentration polarization causes membrane's pores plugging,membrane's fouling, and elevates osmotic pressure differences across themembrane. Critical operation parameters must be predicted to properlyevaluate pressure-driven membranes including the: (1) actual rejectionrate of ions (R_(a)) and concentrations of ions at the membrane (C_(m));(2) saturation degrees of scale pairing ions at the membrane surface andin the reject stream; and (3) actual osmotic pressure differences acrossthe membrane between ions concentrations at the membrane surface(Π_(m)), rather than that in the bulk feed stream (Π_(F)), and that inthe product stream (Π_(P)) [Desalination, 2006, v. 201, pp. 106-113 andpp. 114-120; U.S. Pat. No. 7,093,663].

As shown in FIG. 2, the saturation limit of gypsum within certainconfines increases with increasing sodium chloride concentrations. Eventhough RO membranes reject ions at nearly equal rates, which equallyincrease the background ionic strength (e.g., sodium chloride) that, inturn, increases the solubility limit of gypsum at the membrane surfaceand in the reject brine, the productivity of RO in de-salting seawaterremains limited by two factors: (1) the actual osmotic pressuredifferences across the membrane between the concentrations of rejectedions at the membrane surface and in the product stream (Π_(m)−Π_(P)),thereby brine is reject at a level not exceeding 70,000 mg/L of TDS; and(2) gypsum builds-up at the membranes' surfaces and within membranes'pores wherein imperfect membranes' surfaces and pores act as nucleationsites for gypsum.

In addition to the restrictions on the TDS and gypsum levels, RO (aswell as NF) membranes require extensive conventional pre-treatment toprotect them (e.g., colloidal and suspended solids, carbonate scale,biological growth, etc.). Conventional coagulation-filtration methodswithin the seawater pre-treatment step may be replaced with membranesfiltration such as microfiltration (MF), ultrafiltration (UF), or acombination. For example, FIG. 3 shows a possible general configurationof pressure-driven membranes to pre-treat and de-salt seawater. Screenedand intermittently, not continuously, disinfected (e.g., by chlorine)seawater is mixed with an acid (e.g., sulfuric acid or hydrochloricacid) to lower the pH to about 4, an anti-scale agent, and an anti-foamagent before it is fed to a vacuum de-aerator or an atmosphericstripping tower to strip of carbon dioxide. The de-aerated seawater isthen fed to a filtration step (e.g., MF, UF, or a combination) to removesuspended solids. Depending on the configuration of MF or UF, whetherit's a dead-end or a cross-flow configuration, the reject stream fromthe filtration step may vary between 10-20% of seawater feed stream.Thus, the filtration step may be located upstream, rather thandownstream, of the de-aeration step to reduce the size of the de-aeratoror the stripping tower as well as the amounts of additives. Thede-aerated and filtered seawater may then be de-chlorinated andde-oxygenated by adding an oxygen scavenger (e.g., sodium bisulfite) aswell as alkalinized by a caustic solution (e.g., sodium hydroxide orlime) to at least neutralize the pH before it is de-salted by the RO. NFmay also be integrated as an enabling step before the RO step to protectthe RO membranes from gypsum scale and to presumably increase therecovery ratio of RO.

For seawater with 40,000 mg/L of TDS (e.g., Table 1: S1), RO may besetup in a conventional dual-stage as shown in FIG. 4 (Configuration A)wherein each RO stage is conducted at 25% recovery ratio (RR) and thereject stream from the first stage is used to feed the second stage.About 43% would be recovered from both stages as a total product stream,and about 57% would be rejected as a total reject stream with about70,000 mg/L of TDS. The advantage of such a configuration is that theinherited hydraulic energy within the reject stream from the first stage(e.g., 2.8 bars less than the applied pressure at the seawater feedstream) is directly utilized to feed the second stage. On the otherhand, one of the disadvantages of such a configuration is that the totalproduct stream may not meet the permissible: (1) TDS range in drinkingwater (200-500 mg/L) since the product stream of RO membranes isgenerally entrained with dissolved ions (high but still incomplete ionsrejection) whereas the second stage, in particular, is further strainedby the higher TDS in the reject stream from the first stage (the feedstream for the second stage); and/or (2) stringent boron content indrinking water (0.5 mg/L) since RO membranes are incapable ofefficiently removing boron species at an acidic pH level (boric acidform) or even a near neutral pH range (nearly equally boric acid andborate forms) unless a caustic solution is added to the RO feed streamat a pH level of about 10.5-11 (borate form), which is prohibitive(precipitate magnesium hydroxide). Partial alkalinization of RO feedstream (e.g., sodium hydroxide or lime) is also intricate since it may:(1) violate the permissible TDS range in the product stream (adding moreions) and/or elevate gypsum concentration, thereby promote gypsumsaturation at the membranes' surfaces (e.g., lime); (2) not meet thepermissible boron level in the product stream; and (3) require pHneutralization of both the total RO product stream and reject stream(e.g., by carbon dioxide).

On the other hand, NF membranes allow most of monovalent ions (chloride,sodium, and potassium) to pass through the membranes, partially rejectdivalent cations (magnesium and calcium), but nearly completely rejectsulfate. The low NF rejection of monovalent ions (about 10%) reliefs theosmotic pressure restriction but operating NF within or below thesaturation limit of gypsum at the membrane surface remains the limitingconstrain. For seawater with 40,000 mg/L of TDS (e.g., Table 1: S1),FIG. 4 (Configuration A) also shows that NF can be setup in aconventional dual-stage wherein each stage is conducted at 50% recoveryratio (RR) and the reject stream from the first stage is used to feed tothe second stage. The total product stream from both NF stages would beabout 75% and nearly de-sulfated with about 18% less TDS than seawaterfeed stream. The total reject stream from both stages would be about 25%of the feed stream but heavily sulfate infested. As such, NF may be usedas an enabling step for RO to relief RO from contending with the gypsumsaturation issue but marginally mitigate the RO osmotic pressureconstrain.

If NF is conducted in a dual-stage setup at 75% overall recovery ratioas an enabling step for RO as shown in FIG. 4 (Configuration B), theoverall recovery ratio of RO in a dual-stage setup (reject stream fromthe first stage feeds the second stage) would be about 55%. This meansthat the total RO product stream is about 55% and the total RO rejectstream is about 45% at the TDS limit of about 70,000 mg/L based on theRO feed stream (not the pre-treated seawater feed stream). However, the12% gain in the RO's overall recovery ratio would come at a significantcapital expenditure (the cost of adding the NF step is nearly as much asthe cost of the RO step), an enormous pumping power since both NF and ROare high-pressure membranes, and the combined reject stream (59%) fromboth the NF and RO setups actually exceeds the total reject stream (57%)from the standalone RO setup (FIG. 4, Configuration A).

However, the inventor's suggestion is that an alternative and innovativeRO setup may be envisioned as shown in FIG. 4 (Configuration C) to: (1)avoid the intricacy and expenses of adding and operating NF as anenabling step for RO; (2) resolve the permissible TDS and boron levelsin the RO product stream for drinking water; (3) minimize and confinethe use of a caustic solution only within the RO product; and (4) uselow-pressure RO membranes in the second stage rather than high-pressureRO membranes in both stages. For seawater with 40,000 mg/L of TDS (e.g.,Table 1: S1), RO can thus be setup in a dual-stage wherein the firststage is conducted at 43% recovery ratio (RR) and the product streamfrom the first stage is split into two slip streams (20% and 80%). The80% slip stream of the product stream from the first stage is mixed witha caustic solution and fed to the second RO stage, which is alow-pressure stage (rather than a high-pressure stage as the case in thefirst RO stage) conducted at 90% recovery ratio. The product stream fromthe second stage is then blend with the 20% slip stream of the productstream from the first stage to generate the total RO product stream(about 40% overall recovery). The product stream can be neutralized withdosing carbon dioxide from, for example, the de-aerator. The rejectstream from the second RO stage, which is 10% of the feed stream of thesecond stage (about 3% of the pre-treated seawater feed stream) isrecycled for blending with the pre-treated seawater feed stream to: (1)compensate for the lost 10% as a reject stream from the second RO stage,and therefore to maintain the overall recovery ratio of the productstream at 40%; and (2) slightly reduce the TDS (mitigate the imposedosmotic pressure and gypsum saturation limits on RO) in the pre-treatedseawater feed stream before it is fed to the first stage; and (3)partially alkalinize the pre-treated seawater feed stream before it isfed to the first stage to improve boron removal. The reject stream fromthe first stage remains at 57% with about 70,000 mg/L of TDS as theconventional RO design (FIG. 4, Configuration A), does not require pHneutralization, and its inherited hydraulic energy may be recycled viaan energy recovery device (not shown in FIG. 4, Configuration C) to thepre-treated feed stream.

RO may be the most common de-salting method. However, MSF as athermal-driven de-salting concept produces over 80% of all de-saltedwater in the world. The dominance of MSF is contributed to severalfactors. First, boiling occurs when the vapor pressure of water is equalto the total pressure on the water surface. FIG. 5 shows the boilingpoints of pure water and water containing sodium chloride (anapproximation to a saline stream or a concentrated saline stream) atdifferent concentrations as a function of total pressures. Underatmospheric pressure (1.01 bar), pure water boils at 100° C. whereaswater saturated with sodium chloride boils at 109.5° C. On the otherhand, pure water boils at 44° C. and water saturated with sodiumchloride boils at 50.3° C. under a total pressure of 0.1 bar absolute(sub-atmospheric pressure). MSF (as well as multi-effect distillation,ME) are based on a series of flashing stages; each flashing stagepossesses a lower pressure to lower the boiling point of seawater thanthe previous stage. This allows successive reduction of the boilingpoint of seawater as it gets more concentrated in going down theflashing stages. Such methods are thus based on a multiple boilingconcept under reduced (sub-atmospheric) pressures without supplyingadditional heat after the first flashing stage. Pairing MSF with powerin a power-water desalination co-generation plant to deliberately divertsteam and/or direct exhausted steam from turbines via a brine heater asa heat source for MSF (or ME) dominates seawater de-salting.

Second, a product stream from any thermal-driven de-salting method intreating seawater and the like that contains non-volatile ions is nearlypure distillate. Entrainment of dissolved ions in the product stream (asis the case with RO) is far less pronounced, and thereforethermal-driven de-salting methods produce distillate below thepermissible limits of TDS and boron in drinking water. However,distillate may require blending with brackish water, or lime and carbondioxide, to adjust the TDS (makes it acceptable taste-wise to consumers)and prevent corrosion in distribution pipelines.

Third, MSF or any thermal-driven de-salting method, which is in contrastto pressure-driven methods, is not limited by the osmotic pressure ofseawater and the like. Therefore, it is potentially capable of producingmore distillate and rejecting brine at a level that may reach 250,000mg/L of TDS. However, the prevailing hydrates of calcium sulfate aboveboiling point are hemihydrate and anhydrite (FIG. 2). Such hydrates areless soluble than the sparingly soluble gypsum and their solubilitiesare inversely and drastically decreased with increasing temperaturesabove the boiling point. De-salting seawater at, or above, the boilingpoint would be hindered since such hydrates are the dominant forms ofcalcium sulfate scale. Consequently, for seawater with 40,000 mg/L ofTDS (e.g., Table 1: S1), brine from thermal-driven de-salting methods isrejected at a level not exceeding 65,000 mg/L of TDS as an upper limit.

Once-Through MSF (OT-MSF) is the simplest MSF desalination plant. AnOT-MSF desalination train is depicted in FIG. 6 (Configuration A). Itshould be noted that an MSF desalination plant typically comprisesmultiple standalone OT-MSF trains. The only connections between OT-MSFtrains in a desalination plant may be a seawater intake line to feed theplant and a line to reject brine back to a sea. Seawater feed stream(e.g., screened and chlorinated) is mixed with an acid (to convertbicarbonate to carbon dioxide) along with anti-scale and anti-foam, andpassed through a vacuum de-aerator to remove carbon dioxide. Thepre-treated seawater is then mixed with a caustic solution and an oxygenscavenger and introduced to the last stage of the train (e.g., eachtrain typically comprises 23 heat recovery flashing stages). The brineheater is typically driven by turbines' exhausted steam (e.g., lowpressure steam at 100° C.) to heat the pre-treated seawater in the firstflashing stage to about 90° C. as at a top brine temperature where thepressure is slightly reduced so that it is just below the vaporsaturation pressure of water. Such a relatively low top brinetemperature allows de-salting seawater below the thresholds of theanhydrite scale envelop (FIG. 2). The sudden introduction of heatedseawater into a lower pressure stage causes it to boil so rapidly as toflash into vapor. A relatively small portion of seawater feed stream inthe first flashing stage is converted to vapor. The slightlyconcentrated seawater (brine) in the first stage then passes through therest of the flashing stages where each stage is conducted at a reducedpressure to lower the boiling point of the brine than the previousstage. This allows successive reduction of the boiling point of thebrine as it gets more concentrated in going down the flashing stages andwithout pumping aid. The flashed off vapor condenses on the tubes sideof the condenser/pre-heater units and accumulates across the heatrecovery as distillate. Because the de-aerated seawater feed streamentering the train counter flows with the flashed off brine, thereleased latent heat of the condensed vapor is used to preheat seawateras it enters the last stage of the train and gains more heat as it goesup the flashing stages before it enters the brine heater. The brine fromthe last flashing stage is rejected at about 40° C. and blown down to asea.

In order to operate an OT-MSF desalination train below the anhydritescale envelop, the distillate recovery ratio is forced to be very low(about 10%). At such a low recovery ratio, the TDS gain in reject brineis also low. However, the “Recycle Brine” (RB-MSF) desalination conceptreplaces the OT-MSF desalination concept in the past 20 years to: (1)presumably increase the distillate recovery ratio (e.g., claimed to be30-40%); (2) presumably reduce the volume of seawater feed stream(thereby reducing the size of de-aerators) as well as the volume ofreject brine; and (3) entirely eliminate the de-alkalinization step(adding an acid as a step in seawater pre-treatment) and there-alkalinization step (adding a caustic solution as a step indistillate post-treatment).

FIG. 6 (Configuration B) shows a simplified flow diagram for an RB-MSFdesalination train. The RB-MSF train comprises: (1) a brine heater; (2)a heat recovery section; and (3) a heat rejection section. The RB-MSFtrain differs from the OT-MSF train in several design features. First,the flashing stages are divided into to two sections, typically 20stages for the heat recovery section and 3 stages for the heat rejectionsection, but the total conventional number of flashing stages may remainthe same as in the OT-MSF train (e.g., conventionally 23 stages).

Second, the brine heater is driven by low- and intermediate-pressuresteam to heat recycle brine in the brine heater to about 120° C. beforeits fed to the first flashing stage of the heat recovery section at atop brine temperature of about 110° C. where the pressure is slightlyreduced so that it is just below the vapor saturation pressure of water.A small portion of the recycle brine is flashed off to form vapor in thefirst stage and the remaining and slightly more concentrated recyclebrine (brine) passes through the rest of the flashing stages in the heatrecovery section, wherein each stage is conducted at a reduced pressureto lower the boiling point of the brine than the previous stage, andform more vapor. The formed vapor condenses on the tubes side of thecondenser/pre-heater units and accumulates across the heat recoverysection as distillate, and it is released latent heat is used to preheatthe recycle brine that counter flows to the heat recovery section.

Third, an RB-MSF desalination train is operated at a confinedtemperature range that may extend above the anhydrite scale envelope butbelow the scale envelope of hemihydrate (FIG. 2). At about 120° C., thetime for phase transition between the more soluble metastablehemihydrate and the less soluble stable anhydrite may be much longerthan the detention time in the flashing stages. As such, the operationprinciple of an RB-MSF desalination train relies on manipulating thenature of such calcium sulfate hydrates.

Fourth, brine from the last stage of the heat recovery section passesthrough additional flashing stages in the heat rejection section(pressure is also reduced at each successive stage) to recover morevapor and reduce the temperature of reject brine to about 33° C. Theheat rejection section thus controls the excess heat within brine bydissipating it into a very large volume of cooling seawater stream. Aportion of the pre-heated cooling seawater stream, in which bicarbonatemay thermally breakdown into hydroxide ions and carbon dioxide, is usedas a feed stream and thus is mixed with additives (e.g., anti-scale andanti-foam) and passed through a vacuum de-aerator to strip of carbondioxide. The de-aerated feed stream is mixed with a portion of rejectbrine form the last stage of the heat rejection section to form recyclebrine. Recycle brine is then mixed with an oxygen scavenger andintroduced to the last stage of the heat recovery section. The remaininglarge portion of the pre-heated cooling seawater stream from the heatrejection section is rejected, and combined with reject brine from thefinal stage of the heat rejection before blown down to a sea.

An RB-MSF desalination train is presumed to increase the distillaterecovery ratio to about 30-40%, but when the enormous volume of rejectcooling seawater is considered, the ratio of distillate to totalseawater feed stream (the actual feed for an RB-MSF train to producedistillate and the actual feed for only cooling seawater to be rejected)is about 10%, which is about the same distillate recovery ratio of anOT-MSF train. In addition, an RB-MSF train is more susceptible tohemihydrate and anhydrite scale than an OT-MSF train since it isoperated at a top brine temperature of about 110° C. Furthermore, anRB-MSF train incurs additional enormous operating costs due to at leastthe required high pumping power to circulate and reject an enormousvolume of cooling seawater and to constantly re-circulate an enormousvolume of recycle brine. Yet, reject brine from an RB-MSF train isheavily infested with scale prone species than reject brine from anOT-MSF train, thereby is more harmful not only to marine environment butalso to an RB-MSF desalination plant itself (e.g., alters the naturalions composition of seawater around intake lines). Thus, there is nodifference between an OT-MSF desalination plant and an RB-MSFdesalination plant in terms of actual distillate recovery ratio, but anOT-MSF desalination plant is less damaging to marine environment andmore economic than an RB-MSF desalination plant, and yet both types ofplants remain crippled by sulfate scale.

Produced Water and the Like

Formation water and oil are often concurrently produced as wet oil. Inconventional oil production facilities, higher water cuts in wet oiloccur during the middle or later stage of the primary recovery. Furtherincreases in water cuts also occur during the secondary recovery (e.g.,injection of large amounts of external saline water into hydrocarbonsdeposits) or the tertiary recovery (e.g., injection of steam into heavyoil and bitumen deposits) to sustain, improve, or enhance hydrocarbonsrecovery. For example, FIG. 7 illustrates a conventional wet oilgathering center wherein the bulk of formation water (“oil-in-water”(P/W) stream), which is referred to as produced water, is roughlyseparated from the bulk of oil (“water-in-oil” (W/O) stream) by a wetoil gravity tank.

As water cut rises in wet oil, so do the expenses and the problemsassociated with it. Artificial lift equipment, gathering flow lines, wetoil gathering centers, and produced water treatment and disposal systemsmay reach their operating capacity limits quickly. This forces frequentexpensive modifications and/or expansions to flow lines and centers, ora reduction in wet oil production. Water is thus one of the mostpressing issues in any wet oil production facility.

In contrast to seawater, which is a near uniform stream in term of ionscontent with some exceptions, produced waters are complex and variablestreams (contain varying amounts of oil, gases, ions, and additives).The amounts of produced waters are also substantial, which in recentyears render oil production as a by-product to production of producedwaters. Despite the staggering amounts of produced waters, theirtreatment methods seem to remain conventionally evolving around partialremoval of entrained oil with suspended solids in a standalone, or indirect conjunction with conventional de-salting methods (e.g., RO, NF,electrodialysis, vapor recompression, etc.), transcending the much needde-scaling step.

De-oiling is the first essential step in treating produced waters. Someoil may be dispersed in produced water whereas some of it may bedissolved in produced water. The proportions of dispersed oil anddissolved oil in produced waters vary considerably and depend on thenature and the recovery method of crude oil. The sum of dispersed oiland dissolved oil constitutes the Total Oil Content (TOC).

However, crude oil consists of a very large number of organics, most ofwhich are hydrocarbons and derivatives of hydrocarbons (heteroatoms),and many of which are structurally undetermined or difficult toidentify. Crude oil also contains hydrocarbon gases (e.g., methane,ethane, etc.) and acid gases (e.g., carbon dioxide and hydrogensulfide). Further, crude oil varies in its own mix of organic speciesfor reasons such as the nature, depth and maturity of its deposit andsusceptibility to biodegradation. Classifying organics that may exhibitsomewhat similar properties into a matrix of groups may be useful fortracking organics in crude oil. Crude oil may thus be divided into fivestructural groups: (1) normal and branched paraffins; (2) naphthenes(e.g., monocyclic paraffins, polycyclic paraffins, and theirderivatives); (3) monocyclic aromatics; (4) polycyclic aromatics; and(5) heteroatoms (e.g., species containing nitrogen, sulfur, and oxygen).The content of heteroatoms is one of the distinguishing factors betweenthe highly desirable light oil (low in heteroatoms) and the lessdesirable heavy oil or bitumen (high in heteroatoms).

Heteroatoms are common in heavy crude oil and bitumen, which render themappreciably acidic. Nitrogen-containing species are derived frombiological sources such as porphyrins and amino acids. They areabundantly present in the form of organics with naphthenic and aromaticrings as basic species (e.g., pyridines and amides), neutral species(e.g., alkylhydroxypyridines), and acidic species (e.g., pyrroles). Theyalso present, but to a lesser degree, in the forms organometallics withtransition metals (e.g., vanadium, nickel, copper and iron) as non-basicporphyrin complexes. Sulfur-containing species may include mercaptans,sulfides, and thiophenes. Mercaptans and sulfides occur as cyclic,acyclic, and naphthenoaromatic species whereas thiophenes have aromaticand polyaromatic core structures. Oxygen-containing species mainlyinclude carboxylic acids, phenols and ketones, and to a lesser extentinclude alcohols, ethers and esters.

In carboxylic acids, a carboxyl group also bears a hydroxyl group, andconsequently their compounds are appreciably acidic. The carbonstructures of carboxylic acids correspond with the carbon structures ofhydrocarbons in crude oil that originates from, and thus they reflectthe prevailing type of the hydrocarbons in the crude oil. As such,carboxylic acids are classified according to the substituent that isbonded to the carboxyl group: (1) aliphatic acids have an alkyl chainbound to the carboxyl group; and (2) aromatic acids have an aryl chainbound to the carboxyl group. A mixture of aliphatic carboxylic acids(aliphatic and condensed cycloaliphatic) and condensed aromaticcarboxylic acids constitutes naphthenic acids (C_(n)H_(2n-x)O₂; where“n” is the carbon number and “x” is the hydrogen deficiency). Naphthenicacids are the most abundant of carboxylic acids, and they arepredominant when crude oil is subjected to biodegradation.Biodegradation alters species' proportions of oxygen-containing organicspecies, and the increase in naphthenic acids results in a correspondingdecrease in the relative proportions of other oxygen-containing organicspecies. Biodegradation of crude oil takes place under both aerobic andanaerobic conditions, and it is usually found predominately in immatureheavy oil and bitumen deposits, oil deposits that subjected to waterinjection (especially sulfate-rich source water) to improve oilrecovery, or heavy oil and bitumen deposits that subjected to steaminjection or other thermal treatments.

The relative susceptibility of crude oil to biodegradation may proceedby preferentially destructively metabolizing normal paraffins, branchedparaffins, monocyclic saturated and aromatic hydrocarbons, polycyclicnaphthenic and aromatic hydrocarbons, and finally heteroatoms.Biodegradation greatly enhances the generation of mixtures of oxygen-,oxygen/sulfur-, and oxygen/nitrogen-containing organic species by: (1)generating naphthenic acids via partial oxidation of pure hydrocarbons;(2) generating sulfoxides via the oxidation of sulfides and thiophenesand/or the sulfonation of phenolic and benzylic O-species; and (3)breaking the aromatic or cyclic ring of nitrogen-containing organicspecies (ring-opening) via intermediate pathways (generation ofnitrogen/oxygen-containing organic species) by diluting, if notdepleting, nitrogen, and adding carboxyl or hydroxyl groups. Thus,biodegradation negatively impacts the economy of oil production in termsof quality (e.g., degraded oil due to a decrease in the amount ofparaffins), processibility (e.g., generation of surface active species,promotion of corrosion, impairment and resistant of catalytichydro-treatment in refining crude oil, and emission of NO_(x) andSO_(x)), and treatability of by-product waters (e.g., produced water,refinery's wastewater, and the like).

Carboxylic acids are amphiphilic species. As such, they have hydrophobictail groups (e.g., long chain hydrocarbons, alkylnaphthalene, alkylbenzene, or polysiloxanes) and hydrophilic ionized or polar head groups(e.g. ionic soaps, alkyl benzene sulfonates, or amino acids). Theirunique amphiphilic molecular structures control their adsorption,emulsification, shape, size and packing in water. When crude oil ismixed with water as in wet oil, the intermolecular bonds of theirhydrophobic and hydrophilic groups are distorted; the hydrophobic tailsare oriented toward the oil phase while the hydrophilic heads areoriented toward the water phase. Thus, they aggregate at the oil-waterinterface with their hydrophilic heads in contact with the water phaseand their hydrophobic tails in the oil phase. They may assemble asinterfacial films between the immiscible liquid phases (oil and water),and yet the increase of their concentrations causes them toself-associate as bi-layers, micelles and vesicles in the bulk(dominant) phase. Carboxylic acids with long aliphatic hydrophobicchains pack tightly at the interface, whereas branched and aromaticcarboxylic acids pack loosely. In addition, increasing the length of thehydrophobic groups decreases their aqueous solubilities but increasestheir interface adsorption affinity and ability to form aggregates. Onthe other hand, their hydrophilic groups are surface active when thecarbon number is greater than 7 (species with carbon number<7 may bemore soluble in water, and thus may not be surface active), which causessevere scaling problems. Thus, the existence of carboxylic acids in wetoil is an overwhelming source for forming both emulsion and scale.

TOC is commonly used to evaluate the efficiency of de-oiling producedwater as well as to monitor the toxicity and carcinogenicity of producedwater especially when its discharged to surface water (e.g.,contamination of seawater by offshore platforms); holding ponds anddisposal wells (e.g., contamination of groundwater); or wastewatertreatment facilities. FIG. 8 shows a simplification of TOC in producedwater as the sum of dispersed and dissolved organics. In the U.S., EPAMethod 1664 is the officially approved method for measuring TOC, whichis defined as: “n-hexane extractable material (HEM: oil and grease) andn-hexane extractable material that is not adsorbed by silica gel(SGT-HEM: non-polar material) in surface and saline waters andindustrial and domestic aqueous wastes”. For offshore operations, thecurrent regulation (EPA Method 1664) is a 30-day volume weighted averageTOC in discharged produced water not exceeding 29 mg/L.

Inorganic scale prone species (constitute a portion of the ions content)in produced water, which may include divalent cations, monovalent anddivalent anions, transition metals and silica, are also critical.Depending on the depth and formation type of oil deposits, producedwaters may generally be a chloride-rich hyper-saline stream that is alsorich in both sodium and calcium, or a sulfate-rich saline stream.Chloride-rich produced waters that are high in calcium are generallyhigh in alkaline earth cations (strontium, barium, and in some casesradium) but depleted of sulfate. Some of the naturally occurringisotopes of strontium (Sr-87) and barium (Ba-130 and Ba-132) areradioactive. In addition, the availability of radium in produced watersuggests that the decay series of radium's isotopes is common, and thussuch water may be radioactive. However, sulfate-rich produced waterstend to contain an appreciable concentration of calcium, minorconcentration of strontium, but may be depleted of barium and radium.

Oil producers may have the tendency to follow a so-called “applicationdrift” where specifications from one method or project “drift” intoanother. The same goes for instrumentations where an instrumentoriginally designed for one kind of measurement “drifts” into another.For example, an instrument designed to measure the particle sizedistribution (PSD) of suspended solids “drifted” to measure entrainedoil droplets with suspended solids.

Consequently, partial de-oiling of produced waters is routinely targetedby two or three steps based on the size of entrained oil droplets. Skimtanks may be used as a first de-oiling step to primarily separate thebulk of oil droplets (e.g. >100 μm) from water. Corrugated plateseparators, hydrocyclones, centrifuges, electrostatic, induced gasflotation without chemical addition, and combinations may be used in asecond de-oiling step to remove smaller oil droplets (e.g., 15-100 μm).Sometimes, skim tanks (driven by a gravity force) may be replaced byhydrocyclones or centrifuges (driven by a centrifugal force) to minimizeretention time. Induced gas flotation with chemical addition, adsorption(e.g., activated carbons, nutshells, manganese dioxide, etc.), membranesfiltration (e.g., MF or UF), extraction (e.g., liquid solvents,polymers, or supercritical fluids), and combinations may be used, butnot always, as a final polishing step to remove finer oil droplets. Asshown in FIG. 7, for example, produced water (O/W stream), may betreated by: (1) three conventional de-oiling steps (e.g., skim tanks,gas flotation, and along with adsorption, filtration, or extraction; andthen (2) direct de-salting (if applied) without de-scaling.

However, none of the above mentioned de-oiling steps, individually orcollectively, may be capable of efficiently removing TOC, and thus asignificant portion of TOC may remain in the treated produced water asemulsion-scale and/or toxicity contributors. In addition, such de-oilingsteps are not ZLD since they produce oily waste stream (e.g., skimtanks, hydrocyclones, centrifuges, MF, and UF), oily stripping streams(e.g., extraction by polymers), and exhausted oily adsorption materials(e.g., activated carbon, nutshell, and manganese dioxide) that require aproper disposal path and/or a further treatment. Produced watertreatment thus remains a dual problem since neither the oil phase issufficiently recovered (de-watered) in a useful form nor is the waterphase sufficiently de-oiled, and yet there are still the questions ofthe: (1) disposal of oily waste streams, oily stripping streams,exhausted adsorption materials, or a combination; (2) environmentalimpact of discharging produced water; and (3) beneficial use of producedwater. As a result, offshore produced water discharge limits may befrequently violated; onshore surface discharge of produced water remainsrestricted; and the beneficial usage of produced water by de-saltingmethods remains hindered since neither efficient nor economic de-saltingmethods can be operated in the absence of efficient de-oiling (as wellas de-scaling).

Despite these issues, it is interesting to observe that the so-called“application drifts” that may directly combine conventional inefficientde-oiling methods with conventional de-salting methods (e.g., RO, NF,electrodialysis, vapor recompression, or a combination that wereoriginally developed for de-salting nearly uniform streams such asseawater) are continually proposed to treat produced water. This wouldseem to be unusual since the technical inferiorities of such systems arerecognizable. Beside the need for efficient de-oiling of the water phaseand de-watering of the oil phase as explained above, Table 1 (S6 or S7)and FIG. 2 show, for example, that such scale-infested produced watersare already saturated with gypsum at ambient temperatures. In terms ofions content, such produced waters differ from seawater in two distinctfactors. First, calcium in seawater is about 40% of sulfate whereascalcium in such produced waters substantially exceeds sulfate (about200% of sulfate). Second, the possible breakdown of salt compounds inseawater does not contain calcium chloride, which is in contrast to thepossible breakdown of salt compounds in produced waters since the excessof calcium in such produced waters may exist as calcium chloride. FIG. 9shows that the presence of calcium chloride depresses the solubility ofgypsum while retaining the same solubility patterns as in the presenceof sodium chloride despite the solvation power of sodium chloride due tothe common ion effect (calcium). Thus, directly de-salting such producedwaters would be heavily impaired if not virtually impossible since anyof the calcium sulfate hydrates may be dominant.

The pitfalls of such “application drifts” are reflected, for example, incombining conventional de-oiling with de-salting methods to treatproduced water (e.g., Table 1: S6 and S7). In such application drifts,produced water may be inefficiently treated by a myriad of typicalde-oiling units (e.g., a skim tank, a flotation unit, and a nutshellunit) and in conjunction with (as depicted in FIG. 3) membranefiltration units (MF or UF), as well as membrane de-salting units (RO,NF, or a combination). From strictly de-oiling and de-scalingstandpoints, such a primitive application “drift” does not offer thecapability demanded nor does match the nature and chemistry of producedwaters. The combined de-oiling units and membrane filtration units,despite of their excessiveness, do not remove dissolved oil, which wouldbe carried over to the de-salting units as foulants. Of equalimportance, the de-salting units may nearly immediately impair due toexcessive scale build-up as a result of the already gypsum saturatedproduced waters and the presence of calcium chloride that depresses thesolubility limit of gypsum, which would be further compounded by theeffect of concentration polarization (FIG. 1) at membranes' surfaces.

FIG. 10 may depict another example of an application “drift” wherein thecritical de-oiling and de-scaling problems are not obviated. Producedwater is roughly de-oiled by a skim tank and a flotation unit before itis mixed with additives (e.g., acid, anti-scale and anti-foam) and fedto a feed heat exchanger where it is pre-heated to near boiling by thehot distillate from an evaporator. The pre-heated produced water (brine)is de-aerated by a steam stripper, and mixed with caustic soda and fedto the evaporator where it is temperature is further raised bylow-pressure steam. The evaporator may be a mechanical vaporrecompression (MVR) system, which basically consists of a bulky vaporbody, a large falling-film heat exchanger and a large mechanical vaporcompressor. The brine is continually circulated through the evaporatorvia the top of the falling-film heat exchanger. Vapor is generated asthe brine falls down the tubes of the falling-film heat exchanger,withdrawn into the compressor, and compressed to above the brine'snormal boiling point. The compressed vapor is fed to the falling-filmheat exchanger to transfer its latent heat to the circulated brine.Here, MVR is designed as a conventional single stage evaporator operatedabove the brine's normal boiling point and may be in conjunction withthe use of sodium sulfate or calcium sulfate as a seeding agent topresumably preferentially precipitate anhydrite along with other scaleprone species in the supersaturated brine away from tubes of thefalling-film heat exchanger.

From strictly de-oiling and de-scaling points of view; let alone otherengineering, metallurgical, economical (capital and operating costs) andenvironmental issues that may be prohibitive; the design is inadequatefor two profound reasons. First, the de-oiling steps are deficient sincethey generate a roughly de-oiled stream that carries over dissolved oilto the feed heat exchanger, steam stripper and MVR system. The carriedover dissolved oil acts as a foulant for heat transfer surfaces andcauses severe foaming problems (e.g., a compressor failure and/or aliquid discharge from vapor heads). Second, produced water is alreadysaturated with calcium sulfate before processing. As evaporationprogresses in the feed heat exchanger and the steam stripper prior tothe evaporator, calcium sulfate scale along with other notorious scaleprone species are concentrated, which would cause scalefouling/plugging, reduce heat transfer efficiency, and elevate theboiling point thereby reducing the temperature driving force for heattransfer. The latter is a critical factor in designing an MVR with lowtemperature driving force above normal boiling. Aside from the feed heatexchanger and steam stripper that are directly subjected to calciumsulfate scaling, the seeding concept within the evaporator to minimizetubes plugging is also ineffective. Hemihydrate is the first form ofcalcium sulfate hydrates to precipitate in the evaporator according tothe rule of “stepwise sequence” of phase transformations (from lessstable to more stable forms) and it is precipitation evolves rapidly andfor a relatively finite time (e.g., extends to several hours) comparedto the detention time elapsing during the circulation of brine throughthe evaporator. Thus, the metastable hemihydrate would continuouslydeposit on the heat transfer tubes even though calcium sulfate isreadily supersaturated in the slurry but the anhydrite stable form maynot be attained quickly enough to minimize tubes plugging.

It should also be recognized that the seeding agent must be selected ofthe same form that deposits during evaporation but even if a selectedform of calcium sulfate was used as a seeding agent, different forms ofcalcium sulfate (hemihydrate and anhydrite) would co-exist and vary withthe conditions in the evaporator. If sodium sulfate was used as aseeding agent, on the other hand, the forms of sodium sulfate would havea temperature-solubility phase diagram [e.g., U.S. Pat. Nos. 8,197,696and 7,501,065] that totally differs from the temperature-solubilityphase diagram of calcium sulfate forms. In addition, the seeding agentmust be dispersed in the evaporator in the form of very fine particles,and the amount of the seeding agent must substantially exceed theconcentration of calcium sulfate in produced water. Thus, the seedingmechanism is very difficult to control since the: (1) seeding agent maybe a mismatch (in terms of type, form, particle size, and combinationsof these factors) even though it is in the form of sulfate; and (2)amount of the seeding agent is considerable. As a consequence, theseeding mechanism: (1) requires a high flow rate to evaporate producedwater in the heat transfer tubes, which may diminish the evaporationefficiency; and (2) is not adoptable in multi-stage flash evaporatorswherein the boiling point of brine is successively reduced by reducingpressure (FIG. 5).

It should be pointed it out that produced water in this invention isreferred to any water produced from the exploration and production ofhydrocarbons (gas, liquid, and combinations) including unconventionalsources such as hydro-fracturing and coal-bed methane. Hydro-fracturingis used to fracture and stimulate hydrocarbons deposits by injecting afracturing fluid that may be potable water mixed with a large number ofadditives. During and after fracturing, organics including hydrocarbonsand additives along with ions and water within formation layers aremobilized and brought to the surface as produced water (also refers toas flow-back water). In coal-beds methane, produced water is generatedby removing water that permeates coal-beds thereby reducing thehydrostatic pressure to free methane from the crystal surfaces ofcoal-beds. In contrast to the conventional production of hydrocarbons,produced waters from hydro-fracturing and coal-beds methane flow inlarge amounts in the exploration stage (during fracturing or removingwater) and in the early stage of production, and then the flow drops ormay cease as the production of hydrocarbons increases.

THE OBJECTIVES OF THE INVENTION

Effective methods, individually or collectively, to de-water oil andde-oil water, de-scale the de-oiled water, and de-salt the de-scaledwater are essential for treating “oil-water” streams. The specificobjectives of this invention are to provide methods to treat “oil-water”streams by simultaneously and effectively de-watering the oil phase andde-oiling the water phase, de-scaling the de-oiled water phase, andde-salting the de-scaled water phase. As such, the vertical integrationof source water treatment can be attained.

SUMMARY OF THE INVENTION

The present invention provides a method for treating oil-water stream.The inventive method comprises separating oil from water by a stage ofhydrophobic membranes to produce a de-watered oil stream and a de-oiledwater stream. The oil-water stream is selected from the group consistingof water-in-oil stream, oil-in-water stream, and combinations thereof.

The method for treating oil-water stream further comprises the steps of:(i) pre-demulsifying the oil-water stream by an acid to deactivatesurface active species and convert reactive ionic species in theoil-water stream to acid gas to produce a pre-demulsified oil-waterstream; and (ii) treating the pre-demulsified oil-water stream by thestage of hydrophobic membranes to produce acid gas, the de-watered oilstream, and the de-oiled water stream. The acid is selected from thegroup consisting of hydrochloric acid, perchloric acid, hypochlorousacid, nitric acid, citric acid, sulfuric acid, sulfonic acid, phosphoricacid, formic acid, acetic acid, propionic acid, butyric acid, pentanoicacid, hexanoic acid, pyruvic acid, lactic acid, caproic acid, oxalicacid, malonic acid, succinic acid, glutaric acid, adipic acid, humicacid, fulvic acid, and combinations thereof. Acid gas comprises carbondioxide, hydrogen sulfide, sulfur dioxide, and combinations thereof. Themethod for treating oil-water stream further comprises the steps of: (i)pre-demulsifying the oil-water stream by an external source of acid gasto deactivate surface active species and convert reactive ionic speciesin the oil-water stream to acid gas to produce the pre-demulsifiedoil-water stream; and (ii) treating the pre-demulsified oil-water streamby the stage of hydrophobic membranes to produce acid gas, thede-watered oil stream, and the de-oiled water stream. Such methods forproducing the de-oiled water stream further comprises pre de-scaling by:(i) mixing the de-oiled water stream with a calcium source andleonardite to form pre-precipitates comprising foulants in apre-precipitator unit; and (ii) pre-filtering the pre-precipitates toproduce a pre de-scaled water stream. The calcium source is selectedfrom the group consisting of dolime, calcium oxide, calcium hydroxide,and combinations thereof. Foulants comprise strontium, barium, radium,naturally occurring radioactive materials (NORM), silica, bromide,boron, transition metals, phosphates, carbonates, sulfides, andcombinations thereof. The pre de-scaled water stream further comprisesde-scaling by: (i) mixing the pre de-scaled water stream with theorganic solvent, and either aluminum hydroxide or iron hydroxide to formprecipitates comprising calcium sulfoaluminate or calcium sulfoferratein a precipitator unit; (ii) recovering at least a portion of theorganic solvent by a gas; and (iii) filtering the precipitates toproduce a de-scaled water stream. The gas is selected from the groupconsisting of nitrogen, air, water vapor, and combinations thereof.

The method for treating oil-water stream further comprises the steps of:(i) pre-demulsifying the oil-water stream by an organic solvent in ananionated form to deactivate surface active species and convert reactiveionic species in the oil-water stream to acid gas to produce thepre-demulsified oil-water stream; and (ii) treating the pre-demulsifiedoil-water stream by the stage of hydrophobic membranes to produce acidgas, the de-watered oil stream, and the de-oiled water stream. Theorganic solvent is selected from the group consisting of isopropylamine,propylamine, dipropylamine, diisopropylamine, ethylamine, diethylamine,methylamine, dimethylamine, ammonia, and combinations thereof. Theorganic solvent in the anionated form is produced by reacting theorganic solvent with the acid. The method producing the de-oiled waterstream further comprises de-scaling by: (i) mixing the de-oiled waterstream with either aluminum hydroxide or iron hydroxide to regeneratethe organic solvent, and to form precipitates comprising calciumsulfoaluminate or calcium sulfoferrate in the precipitator unit; (ii)recovering at least a portion of the organic solvent by the gas; and(iii) filtering the precipitates to produce the de-scaled water stream.The de-scaling method, wherein step (i) further comprises mixing thede-oiled water stream with the calcium source. The recovered organicsolvent further comprises reacting it with the acid to produce theorganic solvent in the anionated form.

The method for treating oil-water stream further comprises the steps of:(i) pre-demulsifying the oil-water stream by an aluminum source or aniron source to deactivate surface active species and convert reactiveionic species in the oil-water stream to acid gas to produce thepre-demulsified oil-water stream; and (ii) treating the pre-demulsifiedoil-water stream by the stage of hydrophobic membranes to produce acidgas, the de-watered oil stream, and the de-oiled water stream. Thealuminum source is selected from the group consisting of consisting ofaluminum chloride, aluminum chlorohydrate, aluminum nitrate, aluminumsulfate, aluminum formate, aluminum acetate, or a combination thereof.The iron source is selected from the group consisting of consisting ofiron chloride, iron chlorohydrate, iron nitrate, iron sulfate, ironformate, iron acetate, or a combination thereof. The method producingthe de-oiled water stream further comprises de-scaling by: (i) mixingthe de-oiled water stream with the organic solvent to form precipitatescomprising calcium sulfoaluminate or calcium sulfoferrate in theprecipitator unit; (ii) recovering at least a portion of the organicsolvent by the gas; and (iii) filtering the precipitates to produce thede-scaled water stream. The de-scaling method, wherein step (i) furthercomprises mixing the de-oiled water stream with the calcium source.

The de-scaled water stream is suitable for applications comprisehydrocarbons production, hydrocarbons recovery, acid gas scrubbing, andcombinations thereof. The de-scaled water stream is applicable for adesalination method; the desalination method is selected from the groupconsisting of multi-effect distillation, thermal vapor recompression,mechanical vapor recompression, freezing, membrane distillation, vacuummembrane distillation, osmotic membrane distillation, reverse osmosis,nanofiltration, forward osmosis, electrodialysis, pervaporation, andcombinations thereof. The de-scaled water stream further comprisesdesalination by (i) feeding the de-scaled water stream to aRecycle-Brine Multi-Stage Flash (RB-MSF) desalination train, wherein theRB-MSF desalination train comprises heat recovery, to produce distillateand de-scaled reject brine; and (ii) mixing at least a portion of thede-scaled reject brine with the de-scaled water stream to form recyclebrine to feed the RB-MSF desalination train to produce distillate andde-scaled reject brine. The de-scaled reject brine is suitable forapplications comprise hydrocarbons production, hydrocarbons recovery,chlor-alkali industries, acid gas scrubbing, production of road de-icingsalts, and combinations thereof.

This invention is not restricted to use in connection with oneparticular application. This invention can be used, in general, tode-water oil, de-oil water, scrub acid gas, and de-scale de-oiled water.Further objects, novel features, and advantages of the present inventionwill be apparent to those skilled in the art upon examining theaccompanying drawings and upon reading the following description of thepreferred embodiments, or may be learned by practice of the invention.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 illustrates the concentration polarization profile inpressure-driven membranes.

FIG. 2 illustrates the saturation limits of calcium sulfate hydrates asa function of sodium chloride concentrations.

FIG. 3 illustrates a possible flow diagram for seawater treatment bycombinations of pressure-driven membranes.

FIG. 4 illustrates different configurations for de-salting seawater bypressure-driven membranes.

FIG. 5 illustrates the boiling points of pure water and water containingsodium chloride versus total pressures.

FIG. 6 illustrates simplified configurations for an OT-MSF desalinationtrain and an RB-MSF desalination train to de-salt seawater.

FIG. 7 illustrates a conventional wet oil gathering center.

FIG. 8 illustrates the Total Oil Content (TOC) in produced water.

FIG. 9 illustrates the saturation limits of gypsum as a function of theconcentrations of sodium chloride and calcium chloride.

FIG. 10 illustrates a flow diagram for de-oiling and de-salting producedwater.

FIG. 11 illustrates the precipitation of magnesium hydroxide and thegeneration of gypsum from treating seawater with dolime.

FIG. 12 illustrates the precipitation of magnesium hydroxide and thegeneration of gypsum from treating an RO reject stream with dolime.

FIG. 13 illustrates the precipitation of magnesium hydroxide and thegeneration of gypsum from treating an RB-MSF reject stream with dolime.

FIG. 14 illustrates the precipitation of magnesium hydroxide and thegeneration of gypsum from treating an NF reject stream with dolime.

FIG. 15 illustrates a possible flow diagram for the inventive methods.

DESCRIPTION OF THE PREFERRED EMBODIMENT The Precipitation Concept

I have previously invented the Liquid-Phase Precipitation (LPP) processfor the separation of ionic species from aqueous streams. LPP is basedon mixing an aqueous stream with a suitable solvent at ambienttemperature and atmospheric pressure to form selective precipitates. Thesuitable solvents are those which have the capability to meet two basiccriteria.

The first criteria is the suitability to precipitate targeted ionicspecies (charged inorganics and organics) from aqueous solutions. Theselected organic solvent must be miscible with the aqueous phase. Ofequal importance, the targeted ionic species must be sparingly solublein the organic solvent. The addition of such a solvent to anionic-aqueous solution leads to the capture of part of the watermolecules and reduces the solubility of ionic species in the water whichform insoluble precipitates. The solubility of the targeted ionicspecies in the organic solvent is a critical factor in achieving thedegree of saturation. Therefore, solubility related factors such asionic charge, ionic radius, and the presence of a suitable anion in theaqueous solution play an important role in affecting and characterizingprecipitates formation.

The second criteria is suitability for overall process design. For easeof recovery, the selected solvent must have favorable physicalproperties such as low boiling point, high vapor pressure, high relativevolatility, and no azeotrope formation with water. From a process designstandpoint, the selected solvent must have low toxicity since traces ofthe organic solvent always remain in the discharge stream. Further, theselected solvent must be chemically stable, compatible, and relativelyinexpensive.

Several organic solvents have been identified for potential use in theLPP process. These solvents are isopropylamine (IPA), ethylamine (EA),propylamine (PA), dipropylamine (DPA), diisopropylamine (DIPA),diethylamine (DEA), and dimethylamine (DMA). However, IPA is thepreferred solvent in the LPP process. The preference of IPA isattributed to its high precipitation ability with different ionicspecies, favorable properties (boiling point: 32.4° C.; vapor pressure:478 mmHg at 20° C.); and low environmental risks.

Nitrogen (N₂) can form compounds with only three covalent bonds to otheratoms. An ammonia molecule contains sp³-hybridized nitrogen atom bondedto three hydrogen atoms. An amine molecule contains sp³-hybridizednitrogen atom bonded to one or more carbon atoms. The nitrogen has oneorbital filled with a pair of unshared valence electrons, which allowsthese solvents to act as bases. Thus, the organic solvents (ammonia andamines) are weak bases that could undergo reversible reactions withwater or acids. However, when such solvents react with an acid, theunshared electrons of the solvent are used to form sigma bond with theacid, which would transform the solvent into an anionated (acidified)form. The reaction of isopropylamine with formic acid, for example,produces isopropylamine formate, wherein isopropylamine is the organicsolvent and formate is the anionated form. Such solvents in anionatedforms act as weak acids. The acids that are found useful in thisinvention to generate such solvents in anionated forms comprisehydrochloric acid, perchloric acid, hypochlorous acid, nitric acid,citric acid, sulfuric acid, sulfonic acid, phosphoric acid, formic acid,acetic acid, propionic acid, butyric acid, pentanoic acid, hexanoicacid, pyruvic acid, lactic acid, caproic acid, oxalic acid, malonicacid, succinic acid, glutaric acid, adipic acid, humic acid, fulvicacid, and combinations thereof. Such solvents can be regenerated fromtheir anionated forms by a hydroxide source.

Improving the performance of LPP is always a target. One of theessential improvements is to minimize, if not eliminate, the use of theorganic solvent. Inorganic additives can alternatively replace organicsolvents or can be used in addition to organic solvents to induceprecipitation of targeted species. The suitable inorganic additives forLPP are those that can form an insoluble inorganic-based compound oftargeted charged species in an aqueous stream. Such inorganic additivesshould preferably be recoverable and recyclable, useable as a usefulby-product, or produced locally from reject or waste streams. Also, suchinorganic additives should not, themselves, constitute pollutants.Several inorganic additives were identified, developed, and tested forLPP.

A second targeted improvement for LPP is to produce controllableprecipitates that are uniformly distributed with high yield andpreferably in submicron sizes. Submicron precipitates are fundamentallystable and form spontaneously if a narrow resistance time distributionis improvised and/or a surface active agent (naturally existing orinduced) sufficiently acts as a dispersant to prevent immediateagglomeration of the newly formed precipitates. Submicron precipitatesare thus dispersed phase with extreme fluxionality. On the other hand,non-spontaneous unstable macro-size precipitates will form if givensufficient time to rest.

The state (stabile, metastabe, or unstable) of given precipitates can beexpressed thermodynamically by the Gibbs free energy relation asfollows:ΔG=ΔH−TΔS  (1)where ΔG is the free energy of precipitates (provided by, for instance,mechanical agitation or other means), ΔH is the enthalpy that representsthe binding energy of the dispersed phase precipitates in water, T isthe temperature, and ΔS is the entropy of the dispersed phaseprecipitates (the state of precipitates disorder). The binding energy(ΔH) can be expressed in terms of the surface tension (τ) and theincrease in the surface area (ΔA) as follows:ΔG=τΔA−TΔS  (2)When the introduced free energy into the aqueous stream exceeds thebinding energy of precipitates, individual precipitates are broken downand redistributed. In addition, when a surface active agent is presentin the aqueous stream as an effective dispersant, τ is reduced and thusthe precipitates binding energy is diminished. Furthermore, part of theintroduced energy may not contribute to precipitates' deflocculating butit dissipates in the aqueous stream in the form of heat which reducesviscosity. All of these factors increase precipitates dispersion ordisorder (positive entropy). As such, the change in the entropy (ΔS)quantitatively defines precipitates dispersion (solvation).

The Compressed-Phase Precipitation (CPP) process was thus developed bythe inventor to achieve sub-micron precipitates in certain applications.CPP is conceptually similar to LPP in which the targeted ionic speciesmust be nearly insoluble in the organic solvent whereas the mothersolvent (water) is miscible with the organic solvent. However, thedifference is that fluids in the CPP process can be subjected topressure and/or temperature manipulations, or fluids modifications toforce unusual thermo-physical properties (e.g., exhibit liquid-likedensity but with higher diffusivity, higher compressibility and lowerviscosity).

The fast diffusion combined with low viscosity of a compressed organicsolvent into an aqueous phase produces faster supersaturation oftargeted ionic species, and their possible precipitation in the desiredand sub-micron and micron sizes. Thus, the precipitate's size, sizedistribution, morphology, and structure can be controlled. Achievingfaster supersaturation would, in turn, minimize the use of the organicsolvent, reduce the size of precipitation vessels (a very shortretention time), and allow the recovery of targeted ionic species in thedesired precipitates shape and distribution.

Several factors could influence the performance of the precipitationprocess. Among such factors are: (1) the chemistry of the aqueous streamalong with the identity and concentration of it is targeted species; and(2) the conditions under which precipitation is induced by mixing withadditives (an inorganic additive, an organic solvent, and combinations)with the aqueous stream.

Dolime

Dolime (MgO—CaO), which is calcined dolomite, may nearly contain equalamounts of magnesia and lime as well as minor amounts of other oxides.The hydration of lime in dolime occurs readily at atmospheric pressurewhereas the hydration of magnesia requires an extended reaction timeand/or high pressure and temperature to completely hydrate. In order toconvert dolime to magnesium and calcium tetrahydroxide(Mg(OH)₂—Ca(OH)₂), the hydration reaction of dolime may be carried outin a pressurized vessel at a temperature of about 150° C. to convertoxides to their respective hydroxides. However, the separation ofmagnesium hydroxide from calcium hydroxide in the hydrated dolime isextremely difficult due to their close affinity to water. On the otherhand, if dolime was hydrated with a suitable saline stream (e.g., astream that is rich with magnesium chloride but strictly free or nearlyfree of sulfate), the recovery of magnesium hydroxide would be nearlydoubled since magnesium hydroxide is recovered from both the hydrateddolime and the stream that contains magnesium chloride, therebymagnesium in the stream is replaced with calcium from dolime.

Magnesium-rich chloride-type natural brine is the preferred salinestream since it contains an appreciable concentration of magnesiumchloride (as well as calcium chloride) and it is free or nearly free ofsulfate. The overall hydration reaction of dolime with magnesium-richchloride-type natural brine may be simplified as follows:MgO—CaO+2H₂O+MgCl₂→2Mg(OH)₂↓+CaCl₂  (3)The produced magnesium hydroxide and calcium chloride (Eq. 3) existtogether in two distinct phases. Magnesium hydroxide is formed asprecipitates and recovered as a direct product and/or subsequentlytransformed to other by-products, while the calcium chloride isdissolved in the spent brine since it is extremely soluble in water(solubility limits: 7,750-9,200 meq./L at 20-30° C.). The spent brinemay be rejected in disposal wells. Since the typically employed brinealso contains a very high concentration of calcium chloride (e.g.,higher than the concentration of magnesium chloride) along with thegenerated calcium chloride from the conversion of lime in dolime,calcium chloride may also be recovered from the spent brine (afterprecipitating and recovering magnesium hydroxide) by: (1) a standaloneevaporation process to concentrate calcium chloride to about 13,890meq./L at 175° C.; or (2) a freezing process to concentrate calciumchloride to about 5,230 meq./L at −55° C.

On the other hand, the concentration of magnesium in, for example,seawater is typically much smaller than that in magnesium-richchloride-type natural brines. In addition, roughly about one-third ofmagnesium in normal or relatively normal seawater is in the form ofsulfate and the remaining two-third is in the form of chloride (e.g.,Table 1: S1). The hydration reaction of dolime with seawater may begiven for both magnesium chloride and magnesium sulfate as follows:MgO—CaO+2H₂O+MgCl₂→2Mg(OH)₂↓+CaCl₂  (4a)MgO—CaO+4H₂O+MgSO₄→2Mg(OH)₂↓+CaSO₄.2H₂O  (4b)or may be simplified as follows:2MgO—CaO+6H₂O+MgCl₂+MgSO₄→4Mg(OH)₂↓+CaCl₂CaSO₄.2H₂O   (4)

I have tested the hydration reaction of dolime with de-carbonatedseawater (e.g., Table 1: S1) to precipitate magnesium hydroxide. Asstated above, the possible breakdown of salt compounds in seawater doesnot contain calcium chloride. However, calcium chloride would begenerated if dolime was used to precipitate magnesium hydroxide, whichwould, in turn, depress the solubility limit of gypsum. As shown in FIG.11, about 50% of magnesium hydroxide is precipitated from seawater (anearly equivalent amount of magnesium hydroxide would also be extractedfrom the hydrated dolime itself) at about the saturation limit ofgypsum; that is at 529 meq./L of sodium chloride background in thepresence of the generated calcium chloride. At the 50% precipitationrate of magnesium hydroxide, the co-precipitation of gypsum, which wouldrender the value of the precipitated magnesium hydroxide useless, isavoided. However, about 73% of magnesium hydroxide is precipitated fromseawater at about the saturation limit of gypsum; that is at 529 meq./Lof sodium chloride background but when the effect of the generatedcalcium chloride on the saturation limit of gypsum is ignored. At 73%precipitation rate of magnesium hydroxide, the spent seawater stillcontains the same concentration of sulfate (65 meq./L) as in seawaterfeed stream but much richer with calcium, and therefore the spentseawater is essentially saturated with gypsum. As such, theprecipitation of magnesium hydroxide may be conducted at a confinedprecipitation range (50-73%) that extends above the saturation envelopeof gypsum in the presence of calcium chloride but below the saturationenvelope of gypsum when the effect of calcium chloride is ignored.Within such a precipitation range of magnesium hydroxide from seawater,the possible contamination with gypsum precipitates may be avoidable(e.g., at 50% precipitation rate of magnesium hydroxide) or at leasttolerable (e.g., at <73% precipitation rate of magnesium hydroxide).

Gypsum co-precipitation with magnesium hydroxide is highly undesirablesince: (1) their separation from each other is difficult and expensive;and (2) their combination as a final product has no market value otherthan a wasteful sludge that may be disposed of in landfills. As such,precipitating higher amounts of magnesium hydroxide in a near pure formfrom seawater without being heavily contaminated with gypsumprecipitates when the conditions are more conducive to gypsumprecipitation is simply not practicable. When a supersaturated mixtureof magnesium hydroxide and gypsum is detained in conventional settlingand thickening vessels to produce a settled slurry and spent seawater,water is no longer flowing within the settling slurry and is alsodepleted of sodium chloride (depresses further the solubility ofgypsum). Gypsum (as well as the other hydrates of calcium sulfate) mayrequire an extended detention time to induce precipitation when theconcentration of calcium and sulfate is at saturation and the salinestream is in motion (not in a stagnant condition). A bulk of gypsumprecipitates would thus contaminate magnesium hydroxide precipitates inthe settling slurry. In addition, when the settled slurry isconventionally de-hydrated by evaporation above 95° C., gypsum willtransform to the less soluble hemihydrate and anhydrite forms atelevated temperatures. Such hydrates would heavily precipitate, causesevere scaling problems in pipes and processing equipment, and evendestroy magnesium hydroxide precipitates (the targeted product).

I have also tested the precipitation of magnesium hydroxide using dolimefrom de-carbonated reject streams of seawater de-salting methodsincluding RO, RB-MSF, and NF. FIGS. 12 and 13 reveal, respectively, theprecipitation rates of magnesium hydroxide from the RO reject stream andRB-MSF reject stream. RO and RB-MSF nearly equally concentrate ions intheir reject streams (the concentration factor for RO is 1.64 and forRB-MSF is 1.77). Consequently, the proportions of magnesium chloride andmagnesium sulfate in RO and RB-MSF reject streams remain preserved asnearly as their proportions in seawater (e.g., Table 1: S1, S2 and S3).However, such reject streams are much richer with sulfate than seawater.As a result, the precipitation of magnesium hydroxide from RO and RB-MSFreject streams may be conducted at a confined precipitation range(27-42%) that extends above the saturation envelope of gypsum in thepresence of the generated calcium chloride but below the saturationenvelope of gypsum when the effect of the generated calcium chloride isignored. Within this confined precipitation range of magnesium hydroxidefrom the RO or RB-MSF reject stream, the possible contamination withgypsum precipitates may be avoidable or at least tolerable.

The NF reject stream was generated in my experiments by conventionallyconducting NF in a dual-stage setup at it is 75% maximum possiblerecovery ratio to treat seawater (FIG. 4, Configuration A). Roughlyabout two-third of magnesium in the NF reject stream is in the form ofsulfate and the remaining one-third is in the form of chloride (Table 1:S4). This is opposite to the proportions of magnesium chloride andmagnesium sulfate in seawater (Table 1: S1), which is attributed to thelow rejection of chloride (about 10%), partial rejection of magnesium(about 55%), and near complete rejection of sulfate (about 97%) by NF.The concentration of sulfate in the NF reject stream is about 4-timeshigher than the concentration of sulfate in seawater. This results, asshown in FIG. 14, in a low precipitation rate of magnesium hydroxide(about 20%) in a near pure form from the NF reject stream (as well as anearly equivalent amount of magnesium hydroxide would be extracted fromthe hydrated dolime) at about the saturation limit of gypsum, which isat 692 meq./L of sodium chloride background and in the presence of thegenerated calcium chloride, to avoid gypsum co-precipitation. The spentNF reject stream at such a low precipitation rate of magnesium hydroxidestill contains the same substantial concentration of sulfate (252meq./L) as the NF reject stream but richer with calcium, andconsequently the spent NF reject stream is saturated with gypsum.However, if magnesium hydroxide and gypsum are allowed to progressivelyco-precipitate to near complete precipitation of magnesium hydroxidefrom the NF reject stream as may be useless products, sulfateconcentrations in the spent NF reject stream are still considerable andmay range between 72 and 35 meq./L, which reflect the difference in thesaturation limit of gypsum at 692 meq./L of sodium chloride when thepresence of calcium chloride at 120 meq./L is considered (35 meq./L ofsulfate) or ignored (72 meq./L of sulfate). In addition, theconcentration of the generated calcium chloride that remains dissolvedin the spent NF reject stream is also insignificant to be considered forco-precipitation (with magnesium hydroxide and gypsum as may be uselessproducts) by a costly thermal-driven process (e.g., evaporation orfreezing).

This inventor [e.g., U.S. Pat. No. 8,197,696] teaches the innovativeutilization of an amine solvent to effectively and selectivelyprecipitate magnesium hydroxide from a saline stream, whether the salinestream is only a chloride-rich type or rich with both chloride andsulfate. On the other hand, the useful utility of dolime in selectivelyrecovering magnesium hydroxide as a valuable product from chloride-typenatural brines that are rich with magnesium chloride but free or nearlyfree of sulfate has been well known and extensively explored in theprior art over the past century [e.g., U.S. Pat. Nos. 3,301,633 and3,366,451]. However, such a useful utility is diminished when dolime isapplied to a super sulfate-rich saline stream such as an NF rejectstream since, as explained above, a very low selective recovery ofmagnesium hydroxide is feasible (about 20%), unless magnesium hydroxideis allowed to progressively co-precipitate with gypsum and calciumchloride in a thermally-driven unit [as claimed in U.S. Pat. No.9,045,351], which would practically produce inseparable sludge that hasno value and may be disposed of as waste. As also explained above, it isworth re-iterating that the co-precipitation of gypsum with magnesiumhydroxide by dolime does not equate, by no means, to the removal ofsulfate from sulfate-rich source water.

For applying dolime to a sulfate-rich stream (e.g., seawater) or a supersulfate-rich saline stream (e.g., reject streams from RO, MSF, NF andthe like), a partial selective recovery of magnesium hydroxide in aconfined precipitation range (50-73% for normal or near normal seawater;26-42% for RO or RB-MSF reject stream; and <20% for NF reject stream)must be sought so that the generated gypsum from the double displacementreaction between lime in dolime and magnesium sulfate in the salinestream would be at least within a confined concentration that may extendabove the saturation envelope of gypsum in the presence of calciumchloride but below the saturation envelope of gypsum when the effect ofcalcium chloride is ignored. If gypsum was allowed to precipitate withmagnesium hydroxide, neither a useful product would be recovered norwould sulfate be sufficiently removed from the saline stream. Attemptsto solve such critical issues have been uniquely unsuccessful. Thus, anynew process, economically competitive, but capable of efficientlyremoving sulfate and devoid of generating any useless waste productswould be of great interest. This invention recognizes such a viableinterest, and thus methods have now been developed wherein such issuesand disadvantages can be obviated by efficiently binding theprecipitation of calcium and sulfate in a useful inorganic compound,without the formation of gypsum and/or the forced co-precipitation ofcalcium chloride, thereby not only recovering valuable inorganicby-products but also effectively de-scaling source water.

The De-Oiling/De-Watering Concept

An oil-water stream (e.g., wet oil), depending on it is water cut andviscosity, may be a “water-in-oil” (W/O) stream (may also refer to as aW/O emulsion) or an “oil-in-water” (O/W) stream (may also refer to as anO/W emulsion). The water cut in an oil-water stream is the ratio of thewater volume to the volume of total produced liquids (water and oil). AW/O stream means oil is the “primary” (e.g., continuous or dominant)phase while water is the “secondary” (e.g., dispersed) phase. On theother hand, an O/W stream means water is the “primary” (e.g.,continuous) phase while oil is the “secondary” (e.g., dispersed) phase.Conventional oil-water separation methods are inefficient, whether thestream is a W/O or an O/W, since they basically break down a given“primary” phase into two “secondary” phases, one is richer and the otherone is poorer in the “secondary” phase of the “primary” phase.Consequently, neither water is recovered as a readily de-oiled stream(e.g., does not meet regulations) nor is oil recovered as a readilyde-watered stream (e.g., does not meet specifications).

As can be seen in FIG. 7, which depicts a conventional wet oil gatheringcenter, for example, wherein wet oil may be a W/O stream or an O/Wstream (depending on the water cut and viscosity), is broken down in awet oil gravity tank to further produce two distinct streams; a W/Ostream and an O/W stream. The W/O stream from the wet oil gravity tankstill requires further de-hydrating (de-watering) and washing/de-salting(the oil de-salter may be a single or a dual stage). The source of saltin a W/O stream from a wet oil gravity tank (e.g., FIG. 7) is thecarried over water content within the oil phase, not the oil itself,since salt is dissolved within the water content. Thus, thespecifications for dry oil are 0.2 v % of water content and 10 pounds ofsalt per thousand barrels of oil (PTB). Such specifications are based onan assumed 15,000 mg/L of salt content dissolved within 0.2 v % of thewater content in oil. However, the salinity of water in wet oil varieswidely that may range between 5,000 and 250,000 mg/L. When the watersalinity in wet oil is high, the water content in oil must be reducedaccordingly to meet the specification of salt content in oil (10 PTB).It should be noted that such specifications for the water content andsalt content in oil may also refer to as Basic Sediment and Water(BS&W). On the other hand, a O/W from a wet oil gravity tank (e.g., FIG.7) also demands de-oiling by extensive and multiple steps. It is worthnoting that in some wet oil gathering centers, two-phase (gas-liquids)separators and gravity tanks may be replaced with three-phase(gas-liquid-liquid) separators.

However, water de-oiling and oil de-watering are synonymous. Therefore,they should be simultaneously targeted by an efficient method, ratherthan by an elaborate wet oil gathering center with multiple and costlyinefficient steps that often meet neither produced water regulations nordry oil specifications.

By convention, the term “de-watering” refers to the separation of thewater content from oil, thereby separating the dissolved salt contentwithin the water content from oil. The term “de-oiling” refers to theseparation of the oil content (including all organics) from water;organics that may be found in: (1) crude oil, shale oil, coal oil,bitumen, tar, heating oil, bunker oil, kerosene, diesel fuel, aviationfuel, gasoline, naphtha, synthetic oil, lubricating oil, used or spentmotors oil, waxes, and lubricating greases; (2) refineries andindustrial aqueous wastes such as, for example, sour waters; aromaticsresulting from the cracking of hydrocarbon gases; phenols, amines (e.g.,anilines) and their toxic ligands; benzene polycarboxylic acids (e.g.,benzoic, phthalic, isophthalic, terephthalic, hemimellitic, trimelitic,trimesic, mellophanic, prehnitic, pyromellitic, benzene-pentacarboxylic,and mellitic acids), and the like; (3) vegetable, animal and fish oilsuch as carboxylic acids, saturated or unsaturated; and (4) the like.

Examples of oil-water streams may include, but not limited to: wet oiltwo-phase and/or three-phase separators; slope separators; wet oilgravity tanks; tail waters from de-hydrating oil, washing oil,de-salting oil, and combinations thereof; oil spills and/or dischargesinto surface water (e.g., seawater), groundwater and holding ponds fromoffshore and onshore platforms, offshore and onshore oil pipelines, oilshipping platforms, oil tankers, oil feedstock in power generationplants, and the like; produced water; deficient effluent streams fromproduced water treatment facilities, oily waste streams and oilystripping streams resulting from any conventional produced waterde-oiling methods such as gravity-driven units (e.g., skim tanks and thelike), centrifugal-driven units (e.g., hydrocyclones, centrifuges, andcombinations thereof), filtration units (e.g., flotation,microfiltration, ultrafiltration, and combinations thereof), adsorptionunits (e.g., activated carbons, nutshells, manganese dioxide, andcombinations thereof), and extraction units (e.g., micro-porouspolymers, liquid solvents, supercritical fluids, and combinationsthereof); oily aqueous streams resulting from oil processing andrefining; oily aqueous streams resulting from chemical processing andtreating; oily aqueous streams resulting from processing, recovering andtreating vegetable, animal oil and fish oil; downhole wet oil; and thelike.

The natural demulsification of oil-water starts in some oil reservoirswhere oil might preferentially squeeze through the narrow pores oforganically surface coated rocks (e.g., oil wet sandstone, limestone,dolomite, and combinations thereof) and trapped by impermeable rocks(e.g., clay or shale). In such a natural downhole capillary flow, noshear or differential velocity (velocity is in the direction of theflow) or oil droplets rotation are induced. Thus, capillary flow,especially with low capillary forces, is the most efficient method toseparate oil from water.

My de-oiling/de-watering concept [U.S. Pat. No. 6,365,051 (filed on Oct.12, 1999); U.S. Pat. No. 7,789,159 (filed on Jun. 28, 2008); U.S. Pat.No. 7,934,551 (filed on Feb. 7, 2009); U.S. Pat. No. 7,963,338 (filed onFeb. 27, 2009); and U.S. Pat. No. 8,915,301 (filed on Apr. 26, 2011)] isanalogous to the natural demulsification phenomenon of oil inreservoirs. The inventive concept utilizes the hydrophobic interactionsbetween oil and water as immiscible fluids and a properly configuredhydrophobic membrane would efficiently repel water (the non-wettingfluid) and allow oil (the membrane wetting fluid) to permeate throughthe hydrophobic membrane by applying a low pressure.

Hydrophobic interactions are thermodynamic phase and energy relatedphenomena. The Gibbs free energy, as given in Eq. (1), represents theenergy of interactions between water and hydrophobic molecules. Themixing degree of water and hydrophobic molecules depends largely on theenthalpy, which may be re-expressed as follows:ΔH=2ΔH _(w-h)−Δ_(w-w) −H _(h-h)  (5)where “w” is a water molecule and “h” is a hydrophobic molecule. Waterand hydrophobic molecules would not mix if the water molecule andhydrophobic molecule made more favorable interactions with themselves(“w-w” and “h-h”) than they would make with one another (“w-h”). On theother hand, mixing according to Eq. (1) would be favored by the entropy(the disordering property) and the mixing tendency would increase withtemperature. However, in the absence of a hydrophobic molecule, thegeometry of a polar water molecule in a pure aqueous phase istetrahedron wherein the center of the water molecule is positioned in 6possible hydrogen bonding configurations. When a water molecule in anaqueous phase is replaced by a neutral hydrophobic molecule that may notform a hydrogen bond, one of the edges of the tetrahedron water moleculecollapses, thereby reducing the number of possible hydrogen bondingconfigurations to 3 (instead of 6). This, in turn, cuts the entropy ofthe central water molecule by 50%. Hydrophobic molecules aggregatetogether to minimize the hydrophobic surface interface exposed to watermolecules, and thus the entropy may be expressed as follows:ΔS=S _(w) −S _(h)  (6)where S_(w) is the entropy in the water phase, and S_(h) is the entropyon the hydrophobic surface interface. Eq. (6) implies that the lesshydrophobic surface interface interacts with water, the higher theentropy (favors de-mixing), and thus the lower the Gibbs free energy(Eq. (1)). Therefore, hydrophobic interactions are athermodynamic-driven process that seeks to minimize the free energy byminimizing the mixing between water and hydrophobic molecules.

Hydrophobic membranes are not based on size- or charge-exclusion such ashydrophilic filtration membranes wherein the membranes allow water topass through and reject species based on their sizes (MF, UF, and RO) orcharges (NF). In contrast, hydrophobic membranes do not permit passageof water through the membrane until the water capillary pressure (p_(c))of the hydrophobic membrane is exceeded. p_(c) depends on theinterfacial tension, contact angle, and the pore size distribution ofthe hydrophobic membrane as reflected by the following relation:

$\begin{matrix}{p_{c} = \frac{2\tau_{w - o}\cos\;\theta_{w - o}}{r}} & (7)\end{matrix}$where τ_(w-o) is the water-oil interfacial tension, θ_(w-o) is thecontact angle of a water droplet on the membrane surface in the presenceof oil, r is the radius of the membrane pore. The value of the θ_(w-o)can be related to various interfacial tensions as follows:

$\begin{matrix}{{\cos\;\theta_{w - o}} = \frac{\tau_{m - w} - \tau_{m - o}}{\tau_{w - o}}} & (8)\end{matrix}$where τ_(m-w) is interfacial tension of a membrane in contact withwater, and τ_(m-o) is the interfacial tension of the membrane in contactwith oil. When τ_(m-w) is greater than τ_(m-o) the membrane ishydrophobic (0<θ_(w-o)<90°), which means that the value of p_(c) ispositive and thus the membrane is oil wet that permits the passage ofoil and repels water. However, when τ_(m-w) is lower than τ_(m-o), themembrane is hydrophilic (θ_(w-o)>90°). This means that the value ofp_(c) is negative, and the membrane is water wet that permits thepassage of water and prevents oil from entering the membrane poresagainst the applied pressure (p_(a)).

My de-oiling/de-watering concept using hydrophobic membranes is equallyapplicable for separating organics from each other in organic-organic(non-aqueous) mixtures when the targeted organics in the mixture are notmiscible with each other, and differ in their wettability of hydrophobicmembranes.

Vertical Integration of Source Water Treatment

De-Oiling/De-Watering

I have tested the demulsification of oil-water streams from selected keylocations in a wet oil gathering center by a stage of hydrophobicmembranes. For each of the tested oil-water stream, the hydrophobicmembranes simultaneously produced a de-oiled water stream and ade-watered oil stream. The tested oil-water streams included wet oilfrom a two-phase low-pressure separator with: (1) 18%, 33% and 49% watercuts, which were W/O streams; and (2) 82% water cut, which was an O/Wstream. Additionally, an O/W stream from a wet oil gravity tank with aTOC value of 1,210 mg/L was tested. The values of TOC (as well as TPHand non-TPH) in the de-oiled water streams from the hydrophobicmembranes were measured using EPA Method 1664. The values of the watercontent in the de-watered oil streams from the hydrophobic membraneswere measured using coulometric titration (ASTM D6304).

Table 2 reveals that for all of the tested oil-water streams (W/O andO/W streams), the TOC values in the de-oiled water streams from thestage of hydrophobic membranes were consistently below the EPAregulations (TOC: 42 mg/L as the monthly maximum and 29 mg/L as themonthly average). Similarly, the values of the water content (v %) inde-watered oil streams from the hydrophobic membranes were consistentlyvery low. The salt content in oil, which is based on the water contentin oil, can be derived as follows:C _(S-O)=0.35C _(S-W) SG _(W) v _(r-WO)  (9)where C_(S-O) is the salt content in oil (PTB), C_(S-W) is the saltcontent in water (mg/L), SG_(w) is the specific gravity of water, andv_(r-WO) is the volume fraction of the water content in oil. The valuesof the salt content in water (C_(S-W)) were accurately measured by anion chromatography. Even though the measured values of C_(S-W) were high(about 120,000 mg/L), the values of the salt content in oil for all ofthe tested oil-water streams were consistently and unexpectedly wellbelow 5 PTB (<1.75 PTB), which is attributed to the effectivede-watering of oil by the hydrophobic membranes (very low water contentin oil).

Accordingly, reference is now made to FIG. 15 that shows anoversimplified flow diagram for demulsifying an oil-water stream by astage of hydrophobic membranes. An oil-water stream [1] is fed to thestage of hydrophobic membranes [2], wherein such a stage may comprise aplurality of trains equipped with hydrophobic membranes, and the trainsmay arrange in series, parallel, and combinations thereof. The stage ofhydrophobic membranes [2] separate oil from water to produce ade-watered oil stream [3] and a de-oiled water stream [4]. The oil-waterstream is selected from the group consisting of “water-in-oil” (W/O)stream, “oil-in-water” (O/W) stream, and combinations thereof.

As revealed in Table 2, the hydrophobic membranes have a distinctadvantage in demulsifying oil-water streams, whether the stream is a“water-in-oil” (W/O) or an “oil-in-water” (O/W), which could not beachieved by any known method, not even by may be an elaborate wet oilgathering center (e.g., FIG. 7). For example, the demulsification ofoil-water streams by hydrophobic membranes as a single processing stepmay replace most of the processing steps in a wet oil gathering centerincluding the step of wet oil gravity tanks, the multiple steps ofde-hydrating and washing/de-salting oil, and the multiple steps ofde-oiling produced water. The flexibility and modularity of thehydrophobic membranes also allow ease of capacity additions as normallywater cut rises in wet oil (e.g., aged wells, heavily extracted wells,etc.) compared to conventional wet oil gathering centers that deficientin nature, produce copious amounts of different types of waste, andcombine different extensive methods and bulky equipment (e.g., multipleoperational problems such as not meeting specifications and regulations,high maintenance costs, and lack of flexibility). Additionally,hydrophobic membranes could serve irreplaceable functions in efficientlyseparating oil-water streams in less established, remote, near abandonedor new fields, particularly in multiple low production wet oil wells, inwhich wet oil gathering centers may not be available or justifiable(costs prohibitive).

Although demulsification is usually attained by the stage of hydrophobicmembranes as shown in FIG. 15 (steps [1 to 4]), the use of acid,acid-base, and base interactions as a pre-demulsification step inconjunction with the stage of hydrophobic membranes may be useful insome cases of oil-water streams to interconvert non-neutral (acidic andbasic) organics from either ionic water-soluble forms to non-ionicorganic-soluble forms or vice versa by controlling the pH. Thus, naturalsurface active species such as oxygen-containing organic species (e.g.,carboxylic anions) and their precursors (e.g., surfactants) that may beprevalent in some oil-water streams may be controlled by driving theminto either an organics (oil) phase or a water phase by changing the pH.

In another embodiment, as also shown in FIG. 15 [1, 2, 3, 4, 5A and 6],an oil-water stream [1] is pre-demulsified by an acid [5A] and fed to astage of hydrophobic membranes [2]. The acidification [5A] of theoil-water stream [1] deactivates surface active species (e.g., transformcarboxylic anions into carboxylic acids and the salts of phenols intophenols) as well as converts reactive ionic species to acid gas (e.g.,convert carbonates to carbon dioxide and sulfides to hydrogen sulfide).The deactivation of surface active species and the conversion ofreactive ionic species to acid gas by the acid [5A] before entering thestage of hydrophobic membranes [2] enhances de-mixing by minimizing thefree energy between hydrophobic species (oil) and water. The acid [5A]is selected from the group consisting of hydrochloric acid, perchloricacid, hypochlorous acid, nitric acid, citric acid, sulfuric acid,sulfonic acid, phosphoric acid, formic acid, acetic acid, propionicacid, butyric acid, pentanoic acid, hexanoic acid, pyruvic acid, lacticacid, caproic acid, oxalic acid, malonic acid, succinic acid, glutaricacid, adipic acid, humic acid, fulvic acid, and combinations thereof.The stage of hydrophobic membranes [2] separates oil and acid gas fromwater to produce a de-watered oil stream [3], a de-oiled water stream[4] and acid gas [6]. The acid gas [6] from the stage of hydrophobicmembranes [2] can be utilized to pre-demulsify the oil-water stream [1],thereby minimizing the use of the acid [5A] or may be entirely replacingthe acid [5A] as the production of acid gas [6] may sufficientlyprogress during processing. Acid gas comprises carbon dioxide, hydrogensulfide, sulfur dioxide, and combinations thereof.

In yet another embodiment, external sources of acid gas such aspolluting sources comprise, for example, stacks, steam injectionfacilities for hydrocarbons recovery, and combinations may be used inthis invention. Accordingly, as also shown in FIG. 15 [1, 2, 3, 4, 6Aand 6], an oil-water stream [1] is pre-demulsified by an external sourceof acid gas [6A] and fed to a stage of hydrophobic membranes [2]. Thestage of hydrophobic membranes [2] separates oil and acid gas (the addedexternal source of acid gas as well as acid gas resulting fromconverting reactive ionic species in the oil-water stream [1]) fromwater to produce a de-watered oil stream [3], a de-oiled water stream[4] and acid gas [6]. Here, the dual benefits are the utilization,thereby the containment as well as the replacement of using an acid[5A], of an external polluting source (acid gas) [6A] to pre-demulsifythe oil-water stream [1], and yet the simultaneous utilization of theoil-water stream [1] as a scrubbing fluid for the external source ofacid gas [6A].

In yet another embodiment, as also shown in FIG. 15 [1, 2, 3, 4, 5B and6], an oil-water stream [1] is pre-demulsified by an organic solvent inan anionated form [5B] and fed to a stage of hydrophobic membranes [2].The acidification of the oil-water stream [1] by the organic solvent inthe anionated form [5B] deactivates surface active species and convertsreactive ionic species to acid gas, thereby enhancing de-mixing byminimizing the free energy between hydrophobic species (oil) and water.The organic solvent is selected from the group consisting ofisopropylamine, propylamine, dipropylamine, diisopropylamine,ethylamine, diethylamine, methylamine, dimethylamine, ammonia, andcombinations thereof. The acid, which can be reacted with the organicsolvent to generate the organic solvent in the anionated form, isselected from the group consisting of hydrochloric acid, perchloricacid, hypochlorous acid, nitric acid, citric acid, sulfuric acid,sulfonic acid, phosphoric acid, formic acid, acetic acid, propionicacid, butyric acid, pentanoic acid, hexanoic acid, pyruvic acid, lacticacid, caproic acid, oxalic acid, malonic acid, succinic acid, glutaricacid, adipic acid, humic acid, fulvic acid, and combinations thereof. Assuch, an oil-water stream [1] is pre-demulsified by the organic solventin the anionated form [5B], and fed to a stage of hydrophobic membranes[2], wherein oil and acid gas are separated from water to produce ade-watered oil stream [3], a de-oiled water stream [4] and acid gas [6].The produced acid gas [6] from the stage of hydrophobic membranes [2]can be utilized to pre-demulsify the oil-water stream [1] to at leastminimize the use of the organic solvent in the anionated form [5B] (ormay be entirely replacing the organic solvent in the anionated form [5B]as the production of acid gas may sufficiently progress duringprocessing).

In yet another embodiment, as also shown in FIG. 15 [1, 2, 3, 4, 5C and6], an oil-water stream [1] is pre-demulsified by either an aluminumsource or an iron source [5C] and fed to a stage of hydrophobicmembranes [2]. The aluminum source is selected from the group consistingof aluminum chloride, aluminum chlorohydrate, aluminum nitrate, aluminumsulfate, aluminum formate, aluminum acetate, and combinations thereof.The iron source is selected from the group consisting of iron chloride,iron chlorohydrate, iron nitrate, iron sulfate, iron formate, ironacetate, or a combination thereof. The acidification of the oil-waterstream [1] by the aluminum source or the iron source [5C] deactivatessurface active species and converts reactive ionic species to acid gas,thereby enhancing de-mixing by minimizing the free energy betweenhydrophobic species (oil) and water. Thus, an oil-water stream [1] ispre-demulsified by the aluminum source or the iron source [5C], and fedto a stage of hydrophobic membranes [2], wherein oil and acid gas areseparated from water to produce a de-watered oil stream [3], a de-oiledwater stream [4] and acid gas [6]. The produced acid gas [6] from thestage of hydrophobic membranes [2] can be utilized to pre-demulsify theoil-water stream [1] to at least minimize the use of the aluminum sourceor the iron source [5C] (or may be entirely replacing the aluminumsource or the iron source [5C] as the production of acid gas maysufficiently progress during processing).

The additional benefit for the inventive pre-demulsification step (by anacid [5A], an external source of acid gas [6A], an organic solvent in ananionated form [5B], or an aluminum or iron source [5C]) is that itallows the stage of hydrophobic membranes [2] to serve as a three-phase(gas-liquid-liquid) separator. Such a vital benefit is critical indrastically providing a very simplified new wet oil gathering centerbased on the pre-demulsification and demulsification (hydrophobicmembranes) inventive steps; a wet oil gathering center that may easilymeet dry oil specifications and produced water regulations (see e.g.,Table 2).

A further benefit, which stems from the inventive methods, is thatconventional two-phase (gas-liquid) or three-phase (gas-liquid-liquid)separators, which are very large vessels, can be retrofitted withhydrophobic membranes to replace their coalescing packs. Thus, elaborateconventional wet oil gathering centers that already exist (e.g., FIG. 7)as well as those being designed may have, by such inventivemodifications, the advantages of eliminating the conventional step ofwet oil gravity tanks, the conventional steps of de-hydrating andwashing/de-salting oil (W/O stream), and the conventional steps ofde-oiling produced water (O/W stream).

Yet, a further benefit, which also stems from the inventive methods, isto innovatively design a new flotation device, based on thepre-demulsification (by an acid [5A], an external source of acid gas[6A], an organic solvent in an anionated form [5B], or an aluminum oriron source [5C]) and/or the demulsification (hydrophobic membranes)inventive steps, to effectively treat oil-water streams.

De-Scaling

In addition to the effective de-oiling of water streams [4] by theinventive methods, de-scaling of de-oiled water streams is equallycritically needed. Inspection of Table 1 (S6 and S7), for example,indicates that the ratio of calcium to magnesium and the ratio ofcalcium to sulfate in such produced waters are relatively high(respectively, over 2 and about 2). Since calcium concentration isnearly double sulfate concentration in such produced waters, calciumwould be an appropriate precipitation sink for sulfate to beprecipitated as a useful layered double hydroxides inorganic compound ifit was supplemented with an appropriate trivalent cation (e.g., Al⁺³ orFe⁺³), along with a source of hydroxides. As explained in details above,calcium would be generated as an undesirable cation if dolime wasconventionally used to precipitate magnesium hydroxide from source watercomprises sulfate, which should be avoided to prevent the contamination(e.g., gypsum co-precipitation) of the sought out product (magnesiumhydroxide). However, this invention totally departs from theconventional purpose of utilizing dolime, by innovatively deliberatelygenerating calcium from reacting dolime with source water comprisessulfate. The deliberately generated calcium along with the naturallypresent calcium in produced waters are utilized to precipitate a usefullayered double hydroxides inorganic compound by bounding sulfate in theform of aluminate or ferrate upon the addition of an aluminum source oriron source in a direct precipitation step. As such, the use of dolimein this invention is to produce neither magnesium hydroxide, nor gypsum,nor calcium chloride, nor combinations of such compounds ([0082];[0083]).

The amount of dolime that should be added to such produced waters isthus not governed by the stoichiometric equivalent of magnesium thatnaturally exists in such produced waters, but rather is governed by theneeded amount of calcium to precipitate layered double hydroxidesinorganic compound, which is in contrast to the conventional use ofdolime. Since dolime may roughly contain equal amounts of calcium andmagnesium, the ratio of calcium to magnesium in produced waters uponmixing with dolime, may roughly remain the same (as in produced waterswithout mixing with dolime). As given in Table 1 (S6 and S7), forexample, magnesium is a minor divalent cation in such produced waters,and it's ionic radius (0.65° A) is smaller than the ionic radius ofcalcium (0.98° A). Based on the inventor's testing, magnesium was indeedhomogenously fitted within the structure of the close packedconfiguration of the produced layered double hydroxides inorganiccompound (calcium sulfoaluminate or calcium sulfoferrate).

In this invention, de-scaling of a de-oiled water stream [4], as shownin FIG. 15, may be conducted in several possible approaches. Onepossible approach is to pre-precipitate foulants, and then precipitatesulfate as calcium sulfoaluminate or calcium sulfoferrate. The de-oiledwater stream [4] may contain foulants, most of which may be minor butsome of which may be notorious scale and/or radioactive prone species.Foulants comprise the back-end alkaline cations (strontium, barium,radium, and radium's decay series), silica, bromide, boron, transitionmetals, phosphates, carbonates, sulfides, and combinations thereof. Theradium's decay series also refers to as Naturally Occurring RadioactiveMaterials (NORM). It may be desirable (if not essential) to effectivelyand selectively remove such foulants in a pre-precipitator unit,especially the back-end alkaline cations.

Thus, in one embodiment of this invention, de-scaling of the de-oiledwater stream [4], as shown in FIG. 15, resulting from steps [1, 2, 3 and4] as explained above, steps [1, 2, 3, 4, 5A and 6] as also explainedabove, and steps [1, 2, 3, 4, 6A and 6] as also explained above, isconducted as follows. The de-oiled water stream [4] is mixed with acalcium source [7] and leonardite [8] and fed to a pre-precipitator unit[9], which is attached to a pre-filter [10]. The calcium source isselected from the group consisting of dolime, calcium oxide, calciumhydroxide, and combinations thereof. Leonardite [8] is an inexpensiveadditive, contains a high content of humic acid and some fulvic acid.The pairing of humic acid (in leonardite) with a hydroxide source boundssome of divalent and trivalent ions. Of special interest is thepreferential binding of strontium (e.g., the prevailing back-endalkaline cation in produced waters, Table 1: S6 and S7) by leonardite inthe presence of a hydroxide source, which would render thepre-precipitation step as a useful step to selectively precipitatestrontium along with other foulants, thereby removing such foulants [11]and producing a pre de-scaled water stream. The pre-precipitator unitmay be designed, for example, as a grit pot with a (pre) filter or aplurality of compact (pre) filters. The pre de-scaled water stream [12]from the pre-filter [10] is then mixed with either aluminum hydroxide oriron hydroxide [13], and an organic solvent [14] in a precipitator unit[15], wherein sulfate is precipitated in the form either calciumsulfoaluminate (upon the addition of aluminum hydroxide) or calciumsulfoferrate (upon the addition of an iron hydroxide). The organicsolvent accelerates precipitation and results in very high levels ofsupersaturation within may be a few seconds, which enormously simplifiesthe design of the precipitation unit [15] in terms of size (a compactmodular design with a very short retention time) and effectiveness (afast precipitation of calcium sulfoaluminate or calcium sulfoferrate).The organic solvent is selected from the group consisting ofisopropylamine, propylamine, dipropylamine, diisopropylamine,ethylamine, diethylamine, methylamine, dimethylamine, ammonia, andcombinations thereof. A gas [16] is fed near the bottom of theprecipitator unit [15] to recover the organic solvent. The gas isselected from the group consisting of nitrogen, air, water vapor, andcombinations thereof. The recovered organic solvent [14A] is recycledfor reuse in the precipitator unit [15]. The outlet stream [17] from theprecipitator unit [15] is fed to a filter [18] to remove theprecipitates [19] and produce a de-scaled water stream [20].

In another embodiment, de-scaling of the de-oiled water stream [4], asshown in FIG. 15, resulting from steps [1, 2, 3, 4, 5B and 6] asexplained above, is conducted as follows. In addition to using theorganic solvent in the anionated form [5B] to pre-demulsify theoil-water stream [1], a further innovative purpose for using the organicsolvent in the anionated form is that the carried over organic solventwith the de-oiled water stream [4] can be regenerated from its anionatedform, and thus can be directly utilized to precipitate sulfate from thede-oiled water stream [4] upon mixing with either aluminum hydroxide oriron hydroxide, and further upon mixing with a calcium source. As such,the processing steps [8 to 12, and 14], as shown in FIG. 15, areeliminated. Accordingly, the de-oiled water stream [4] is mixed witheither aluminum hydroxide or iron hydroxide [13] to regenerate theorganic solvent from it is anionated form, and to form precipitatescomprising sulfate (either calcium sulfoaluminate upon mixing withaluminum hydroxide or calcium sulfoferrate upon mixing with ironhydroxide) in a precipitator unit [15]. A gas [16] is fed near thebottom of the precipitator unit [15] to recover the regenerated organicsolvent [14A]. The gas is selected from the group consisting ofnitrogen, air, water vapor, and combinations thereof. A calcium source[7] may also be mixed with the de-oiled water stream [4], as needed, tobalance calcium concentration in the de-oiled water stream [4]. Thecalcium source is selected from the group consisting of dolime, calciumoxide, calcium hydroxide, and combinations thereof. The recoveredorganic solvent [14A] is reacted with an acid (not shown in FIG. 15) toproduce the organic solvent in the anionated form [5B] for reuse (topre-demulsify the oil-in-water stream [1]). The acid is selected fromthe group consisting of hydrochloric acid, perchloric acid, hypochlorousacid, nitric acid, citric acid, sulfuric acid, sulfonic acid, phosphoricacid, formic acid, acetic acid, propionic acid, butyric acid, pentanoicacid, hexanoic acid, pyruvic acid, lactic acid, caproic acid, oxalicacid, malonic acid, succinic acid, glutaric acid, adipic acid, humicacid, fulvic acid, and combinations thereof. The outlet stream [17] fromthe precipitator unit [15] is fed to a filter [18] to remove theprecipitates [19] and produce a de-scaled water stream [20].

In yet another embodiment, de-scaling of the de-oiled water stream [4],as shown in FIG. 15, resulting from the steps [1, 2, 3, 4, 5C and 6] asexplained above, is conducted as follows. In addition to using eitherthe aluminum source or iron source [5C] to pre-demulsify the oil-waterstream [1], a further innovative purpose for using either the aluminumsource or the iron source is that the carried over trivalent cation(either aluminum or iron) with the de-oiled water stream [4] is alsoutilized to precipitate sulfate upon mixing with an organic solvent as ahydroxide source, and further upon mixing with a calcium source, asneeded, to balance calcium concentration in the de-oiled water stream[4]. The calcium source is selected from the group consisting of dolime,calcium oxide, calcium hydroxide, and combinations thereof. As such, theprocessing steps [8 to 13], as shown in FIG. 15, are eliminated.Accordingly, the de-oiled water stream [4] is mixed with an organicsolvent [14] to form precipitates comprising sulfate (either calciumsulfoaluminate upon mixing with the aluminum source, or calciumsulfoferrate upon mixing with the iron source) in a precipitator unit[15]. The organic solvent is selected from the group consisting ofisopropylamine, propylamine, dipropylamine, diisopropylamine,ethylamine, diethylamine, methylamine, dimethylamine, ammonia, andcombinations thereof. A gas [16] is fed near the bottom of theprecipitator unit [15] to recover the organic solvent. The gas isselected from the group consisting of nitrogen, air, water vapor, andcombinations thereof. The recovered organic solvent [14A] is recycledfor reuse in the precipitator unit [15]. A calcium source [7] may alsobe mixed with the de-oiled water stream [4] to balance, as needed,calcium concentration in the de-oiled water stream [4]. The calciumsource is selected from the group consisting of dolime, calcium oxide,calcium hydroxide, and combinations thereof. The outlet stream [17] fromthe precipitator unit [15] is fed to a filter [18] to remove theprecipitates [19] and produce a de-scaled water stream [20].

The precipitation of calcium sulfoaluminate or calcium sulfoferratetakes place based on the conditions under which it is effectivelyprecipitated. Based on the inventor's testing, the removal of sulfatefrom source water in the form of either calcium sulfoaluminate orcalcium sulfoferrate is consistently over 97%. One structural formulathat may generally describe certain embodiments of calciumsulfoaluminate or calcium sulfoferrate is as follows:└Ca⁺²┘_(A)└SO₄ ⁻²┘_(B)└M⁺³┘_(C)[xH₂O]where A is the stoichiometric amount of calcium (Ca⁺²), B isstoichiometric amount of sulfate (SO₄ ⁻²), C is the stoichiometricamount of the trivalent cation (M⁺³; which is either aluminum: Al⁺³ oriron: Fe⁺³), and x is the hydration content. Depending on the amount ofsulfate in source water, the chemistry of source water, and the basicitycondition under which sulfate is precipitated in the form of eithercalcium sulfoaluminate or calcium sulfoferrate, the stoichiometric ratio(meq./L) of sulfate to calcium (B/A) is 0.2 to 0.5, the stoichiometricratio (meq./L) of sulfate to the trivalent cation (B/C) is 0.5 to 1.5,and the hydration content (x) is 24 to 32.

The de-scaled water stream [20] is readily usable in vital applicationscomprise hydrocarbons production (e.g., hydro-fracturing), hydrocarbonsrecovery (sustain, improve, and enhance), acid gas scrubbing, andcombinations thereof.

De-Salting

The de-scaled water stream [20] is also readily suitable to feed anydesalination method. The desalination method is selected from the groupconsisting of multi-stage flash desalination, multi-effect distillation,thermal vapor recompression, mechanical vapor recompression, freezing,membrane distillation, vacuum membrane distillation, osmotic membranedistillation, reverse osmosis, nanofiltration, forward osmosis,electrodialysis, pervaporation, and combinations thereof.

However, as highlighted above ([0012]-[0016]), RB-MSF desalinationplants are dominant over the past 20 years and produce over 80% of allde-salted water in the world. A typical seawater de-salting plant, forexample, may consist of eight conventional RB-MSF desalination trains,and each train may consist of 23 flashing stages. Each train may thus bedesigned to produce about 15 million gallons per day (about 357,000barrels per day) of distillate. As such, a single RB-MSF desalinationtrain (FIG. 6, Configuration B) may easily meet, if not exceed, thenormal need of distillate in oil-gas fields' applications (e.g.,generating steam, washing/de-salting oil, etc.). The MSF desalinationconcept, in general, and the RB-MSF desalination concept, in particular,is inapplicable in oil-gas fields. One of the detrimental reasons isthat the heat rejection section of an RB-MSF desalination train requiresan enormous amount of cooling water (e.g., may be about 7-times theamount of distillate), which is simply not available in oil-gas fields.A second profound reason is that the distillate recovery ratio would beextremely low due to the harsh chemistry of produced waters (e.g., Table1: S6 and S7; FIG. 2: the hydrates of calcium sulfate), which rendersthe recycle brine (RB) concept totally useless, thereby requiring anenormous amount of feed stream. This also means that the generatedamount of heavily scale-infested reject brine may be at least over9-times the amount of distillate, which is also prohibitive in oil-gasfields (it requires disposal wells that can withstand such a largeamount of a heavily scaled stream, which is very unlikely).

On the other hand, the salt content in the de-scaled water stream [20]by the inventive de-scaling methods is essentially sodium chloride. Thede-scaling methods in this invention facilitate, in turn, improving theRB concept in conjunction with MSF desalination. As shown in FIG. 15,the RB-MSF desalination train [20, 20A, 21, 22, 23, 24 and 24A] in thisinvention distinctly differs from a conventional RB-MSF desalinationtrain (FIG. 6, Configuration B) in that: (1) the heat rejection sectionof the conventional RB-MSF train is entirely eliminated, which entirelyeliminates the need for an enormous amount of cooling water; and (2)brine may be rejected at a level not exceeding 250,000 mg/L of TDS,which maximizes distillate production, minimizes the required amount ofthe de-scaled water stream [20] to feed the train and reduces the amountof reject brine; and (3) reject brine is subsequently de-scaled, whichis readily usable in applications comprise hydrocarbons production,hydrocarbons recovery, chlor-alkali industries, acid gas scrubbing,production of road de-icing salts, and combinations thereof. Thus, theprofound obstacles that preclude the use of a conventional RB-MSFdesalination train in oil-gas fields are obviated by the newly designedRB-MSF desalination train in this invention.

As such, as also shown in FIG. 15, the de-scaled water stream [20] isfed to the last flashing stage of the RB-MSF train [21], thereby passingthrough the rest of the flashing stages before it enters a brine heater[22]. The inventive RB-MSF train [21] is only for heat recovery (no heatrejection section; therefore, no cooling water), and the number of theflashing stages is easily extendable since the use of the de-scaledwater stream [20] lifts the imposed restriction on top brinetemperature. The brine heater [22] heats the feed stream [20] before itenters the first flashing stage of the train [21] to a pre-designed topbrine temperature where the pressure in the first flashing stage isslightly reduced so that it is just below the vapor saturation pressureof water. This sudden introduction of the feed stream [20] into a lowerpressure flashing stage causes water to boil so rapidly as to flash intovapor and to produce slightly concentrated brine. The slightlyconcentrated brine from the first stage then passes through the rest ofthe flashing stages of the train [21] where each stage is conducted at areduced pressure to lower the boiling point of the brine than theprevious stage. This allows successive reduction of the boiling point ofthe brine as it gets more concentrated in going down the flashing stagesof the train and without pumping aid until the brine [24] is rejectedfrom the last stage. The vapors condense on the tubes side ofcondenser/pre-heater units and accumulate across the train [21] asdistillate [23]. Since the train [20] is fed in a counter flow with theflashed off brine, the released latent heat of the condensed vapors isutilized to preheat the feed stream [20] as it enters the last stage ofthe train and gains more heat as it goes up the flashing stages beforeit enters the brine heater [22]. At the beginning of the operation, thereject brine [24] from the last flashing stage of the train [21] iscompletely recycled for blending with the de-scaled water stream [20] toform recycle brine [20A] as a feed for the train. Once a steady state isattained by establishing a desired concentration factor for recyclebrine [20A], which consists of at least a portion of the reject brine[24] and the de-scaled water stream [20], the remaining portion of thereject brine [24A] may be blown down at a level not exceeding 250,000 ofTDS.

TABLE 1 Samples of Source Water. Ion (meq./L) S1 S2 S3 S4 S5 S6 S7 Na⁺529.1 873.5 1,091.4 692.6 1613.8 337.0 59.2 K⁺ 10.7 26.2 18.9 12.1 32.69.7 3.1 Mg⁺² 125.9 191.7 209.6 332.1 384.0 49.4 22.3 Ca⁺² 27 41.9 47.252.5 82.4 117.5 55.1 Sr⁺² 0.2 0.5 0.5 1.6 0.7 Cl⁻ 623 1,020.1 1,181.2823.5 1900.2 462.6 96.5 HCO₃ ⁻ 2.3 4.2 3.4 8.2 4.0 26.6 SO₄ ⁻² 64.6106.2 114.6 251.6 197.0 61.5 28.1 MgCl₂/ 0.68 0.67 0.68 0.38 0.70 Σ MgCa⁺²/SO₄ ⁻² 0.42 0.40 0.41 0.21 0.42 1.91 1.96 Ca⁺²/Mg⁺² 0.21 0.22 0.230.16 0.21 2.38 2.18 Mg⁺²/SO₄ ⁻² 1.95 1.81 1.83 1.32 1.95 0.80 0.80 S1:Seawater; S2: RO reject stream from seawater treatment at 43% overallrecovery ratio; S3: RB-MSF reject brine from seawater treatment; S4: NFreject stream from seawater treatment at 75% overall recovery ratio; S5:reject stream from flue gas de-sulfurization (spent seawater makeup);S6: produced water; S7: produced water; Σ Mg = MgCl₂ + MgSO₄.

TABLE 2 Hydrophobic Membranes: De-Oiled Water and De-Watered OilStreams. De-Oiled Water De-Watered Stream (mg/L) Oil Stream Stream TOCTPH non-TPH (v %) W/O-LPS (wet oil water cut: 18%) 18.5 12.1 6.4 0.004W/O-LPS (wet oil water cut: 33%) 18.3 12.5 5.8 0.002 W/O-LPS (wet oilwater cut: 49%) 20.1 14.2 5.9 0.003 O/W-LPS (wet oil water cut: 82%)15.6 10.3 5.3 0.002 O/W-WOGT 4.1 0.001 W/O: Water-in-Oil; O/W:Oil-in-Water; LPS: a Low-Pressure Separator; WOGT: a Wet Oil GravityTank.

What is claimed is:
 1. A method for treating an oil-water stream, saidmethod comprising separating said oil-water stream by a hydrophobicmembrane to produce a de-watered oil stream, and a de-oiled waterstream; determining that said de-watered oil stream is less than orequal to 10 pounds of salt per thousand barrels of oil (PTB), saidde-oiled water stream is less than or equal to 42 mg/L of total oilcontent (TOC), and combinations thereof; and wherein said oil-waterstream is a water-in-oil (W/O) stream, an oil-in-water (O/W) stream, andcombinations thereof.
 2. A method for treating an oil-water stream, saidmethod comprising demulsifying said oil-water stream with a pollutingacid gas source, followed with separating by a hydrophobic membrane toproduce a de-watered oil stream, a de-oiled water stream, and an acidgas stream; determining that said de-watered oil stream is less than orequal to 10 pounds of salt per thousand barrels of oil (PTB), saidde-oiled water stream is less than or equal to 42 mg/L of total oilcontent (TOC), and combinations thereof; appreciating that said acid gasstream is carbon dioxide, hydrogen sulfide, sulfur dioxide, andcombinations thereof; and recognizing that said oil-water stream is awater-in-oil (W/O) stream, an oil-in-water (O/W) stream, andcombinations thereof.
 3. The method of claim 2, wherein said pollutingacid gas source is selected from the group consisting of emissions frompolluting stacks, steam injection facilities for hydrocarbons recovery,and combinations thereof.
 4. The method of claim 2, comprising the stepof replacing at least a portion of said polluting acid gas source by theproduced said acid gas stream from said hydrophobic membrane.
 5. Themethod of claim 2, further comprises the step of replacing saidpolluting acid gas source or at least a portion of said polluting acidgas source with an acid, wherein said acid is selected from the groupconsisting of hydrochloric acid, perchloric acid, hypochlorous acid,nitric acid, citric acid, sulfuric acid, sulfonic acid, phosphoric acid,formic acid, acetic acid, propionic acid, butyric acid, pentanoic acid,hexanoic acid, pyruvic acid, lactic acid, caproic acid, oxalic acid,malonic acid, succinic acid, glutaric acid, adipic acid, humic acid,fulvic acid, and combinations thereof.
 6. A method for separatingfoulants from a de-oiled water stream, said method comprising the stepsof: (a) mixing said de-oiled water stream with a calcium source andleonardite to form precipitates comprising said foulants in apre-precipitator unit; wherein said foulants comprise strontium, barium,radium, naturally occurring radioactive materials (NORM), silica,bromide, boron, transition metals, phosphates, carbonates, sulfides, andcombinations thereof; and wherein said calcium source is selected fromthe group consisting of dolime, calcium oxide, calcium hydroxide, andcombinations thereof; and (b) removing said precipitates by a filter toproduce a de-fouled water stream.
 7. The method of claim 6, comprisingthe steps of: (a) separating sulfate from said de-fouled water streamby: (i) mixing said de-fouled water stream with an organic solvent, andaluminum hydroxide or iron hydroxide, to form precipitates comprisingcalcium sulfoaluminate or calcium sulfoferrate in a precipitator unit;wherein said organic solvent is selected from the group consisting ofisopropylamine, propylamine, dipropylamine, diisopropylamine,ethylamine, diethylamine, methylamine, dimethylamine, ammonia, andcombinations thereof; (ii) recovering at least a portion of said organicsolvent by introducing an inert gas stream into said precipitation unit,wherein said inert gas stream is selected from the group consisting ofnitrogen, air, water vapor, and combinations thereof; (iii) filteringsaid precipitates to produce a de-scaled water stream; (b) utilizing atleast a portion of said de-scaled water stream for hydrocarbonsproduction, hydrocarbons recovery, acid gas scrubbing, and combinationsthereof; and (c) desalinating at least a portion of said de-scaled waterstream by a desalination method to produce a distillate stream and ade-scaled reject brine stream; wherein said desalination method isselected from the group consisting of multi-stage flash desalination,multi-effect distillation, thermal vapor recompression, mechanical vaporrecompression, freezing, membrane distillation, vacuum membranedistillation, osmotic membrane distillation, reverse osmosis,nanofiltration, forward osmosis, electrodialysis, pervaporation, andcombinations thereof; and wherein said de-scaled reject brine stream isutilized for hydrocarbons production, hydrocarbons recovery,chlor-alkali industries, acid gas scrubbing, production of road de-icingsalts, and combinations thereof.
 8. The method of claim 7, furthercomprises the steps of: (a) feeding at least a portion of said de-scaledwater stream to a Recycle-Brine Multi-Stage Flash (RB-MSF) desalinationtrain, wherein said RB-MSF desalination train includes only a heatrecovery section, to produce said distillate stream and said de-scaledreject brine; and (b) mixing at least a portion of said de-scaled rejectbrine with said de-scaled water stream to form a recycle brine stream,and feeding said recycle brine stream to said RB-MSF desalination trainto produce said distillate stream and said de-scaled reject brine;wherein a portion of said de-scaled reject brine is discharged from saidRB-MSF train at a level not exceeding 250,000 mg/L of total dissolvedsolids (TDS).
 9. A method for treating an oil-water stream, said methodcomprising demulsifying said oil-water stream with an anionated organicsolvent, followed by separating by a hydrophobic membrane to produce ade-watered oil stream, a de-oiled water stream, and an acid gas stream;determining that said de-watered oil stream is less than or equal to 10pounds of salt per thousand barrels of oil (PTB), said de-oiled waterstream is less than or equal to 42 mg/L of total oil content (TOC), andcombinations thereof; appreciating said acid gas stream is carbondioxide, hydrogen sulfide, sulfur dioxide, and combinations thereof; andrecognizing that said oil-water stream is a water-in-oil (W/O) stream,an oil-in-water (O/W) stream, and combinations thereof.
 10. The methodof claim 9, wherein said anionated organic solvent is generated byreacting an organic solvent with acid; wherein said organic solvent isselected from the group consisting of isopropylamine, propylamine,dipropylamine, diisopropylamine, ethylamine, diethylamine, methylamine,dimethylamine, ammonia, and combinations thereof; and wherein said acidis selected from the group consisting of hydrochloric acid, perchloricacid, hypochlorous acid, nitric acid, citric acid, sulfuric acid,sulfonic acid, phosphoric acid, formic acid, acetic acid, propionicacid, butyric acid, pentanoic acid, hexanoic acid, pyruvic acid, lacticacid, caproic acid, oxalic acid, malonic acid, succinic acid, glutaricacid, adipic acid, humic acid, fulvic acid, and combinations thereof.11. The method of claim 9, comprising the steps of: (a) separatingsulfate from said de-oiled water stream by: (i) mixing said de-oiledwater stream with aluminum hydroxide or iron hydroxide to regenerate theorganic solvent from said anionated organic solvent, and to formprecipitates comprising calcium sulfoaluminate or calcium sulfoferratein a precipitator unit; (ii) recovering at least a portion of theregenerated said organic solvent by introducing an inert gas stream intosaid precipitation unit, wherein said inert gas stream is selected fromthe group consisting of nitrogen, air, water vapor, and combinationsthereof; (iii) filtering said precipitates to produce a de-scaled waterstream; (iv) reacting the recovered said organic solvent in step (ii)with acid to produce said anionated organic solvent; (b) utilizing atleast a portion of said de-scaled water stream for hydrocarbonsproduction, hydrocarbons recovery, acid gas scrubbing, and combinationsthereof; and (c) desalinating at least a portion of said de-scaled waterstream by a desalination method to produce a distillate stream and ade-scaled reject brine stream; wherein said desalination method isselected from the group consisting of multi-stage flash desalination,multi-effect distillation, thermal vapor recompression, mechanical vaporrecompression, freezing, membrane distillation, vacuum membranedistillation, osmotic membrane distillation, reverse osmosis,nanofiltration, forward osmosis, electrodialysis, pervaporation, andcombinations thereof; and wherein said de-scaled reject brine stream isutilized for hydrocarbons production, hydrocarbons recovery,chlor-alkali industries, acid gas scrubbing, production of road de-icingsalts, and combinations thereof.
 12. The method of claim 11, whereinstep (a) (i) further comprises the step of mixing said de-oiled waterstream with a calcium source; wherein said calcium source is selectedfrom the group consisting of dolime, calcium oxide, calcium hydroxide,and combinations thereof.
 13. The method of claim 11, further comprisesthe steps of: (a) feeding at least a portion of said de-scaled waterstream to a Recycle-Brine Multi-Stage Flash (RB-MSF) desalination train,wherein said RB-MSF desalination train includes only a heat recoverysection, to produce said distillate stream and said de-scaled rejectbrine; and (b) mixing at least a portion of said de-scaled reject brinewith said de-scaled water stream to form a recycle brine stream, andfeeding said recycle brine stream to said RB-MSF desalination train toproduce said distillate stream and said de-scaled reject brine; whereina portion of said de-scaled reject brine is discharged from said RB-MSFtrain at a level not exceeding 250,000 mg/L of total dissolved solids(TDS).
 14. A method for treating an oil-water stream, said methodcomprising demulsifying said oil-water stream by an aluminum source oran iron source, followed by separating utilizing a hydrophobic membraneto produce a de-watered oil stream, a de-oiled water stream, and an acidgas stream; determining that said de-watered oil stream is less than orequal to 10 pounds of salt per thousand barrels of oil (PTB), saidde-oiled water stream is less than or equal to 42 mg/L of total oilcontent (TOC), and combinations thereof; wherein said acid gas streamcomprises carbon dioxide, hydrogen sulfide, sulfur dioxide, andcombinations thereof; and wherein said oil-water stream is awater-in-oil (W/O) stream, an oil-in-water (O/W) stream, andcombinations thereof.
 15. The method of claim 14, wherein said aluminumsource is selected from the group consisting of aluminum chloride,aluminum chlorohydrate, aluminum nitrate, aluminum sulfate, aluminumformate, aluminum acetate, and combinations thereof; and wherein saidiron source is selected from the group consisting of iron chloride, ironchlorohydrate, iron nitrate, iron sulfate, iron formate, iron acetate,and combinations thereof.
 16. The method of claim 14, comprising thesteps of: (a) separating sulfate from said de-oiled water stream by: (i)mixing said de-oiled water stream with an organic solvent to formprecipitates comprising calcium sulfoaluminate or calcium sulfoferratein a precipitator unit; wherein said organic solvent is selected fromthe group consisting of isopropylamine, propylamine, dipropylamine,diisopropylamine, ethylamine, diethylamine, methylamine, dimethylamine,ammonia, and combinations thereof; (ii) recovering at least a portion ofsaid organic solvent by introducing an inert gas stream into saidprecipitation unit, wherein said inert gas stream is selected from thegroup consisting of nitrogen, air, water vapor, and combinationsthereof; (iii) filtering said precipitates to produce a de-scaled waterstream; (b) utilizing at least a portion of said de-scaled water streamfor hydrocarbons production, hydrocarbons recovery, acid gas scrubbing,and combinations thereof; and (c) desalinating at least a portion ofsaid de-scaled water stream by a desalination method to produce adistillate stream and a de-scaled reject brine stream; wherein saiddesalination method is selected from the group consisting of multi-stageflash desalination, multi-effect distillation, thermal vaporrecompression, mechanical vapor recompression, freezing, membranedistillation, vacuum membrane distillation, osmotic membranedistillation, reverse osmosis, nanofiltration, forward osmosis,electrodialysis, pervaporation, and combinations thereof; and whereinsaid de-scaled reject brine stream is utilized for hydrocarbonsproduction, hydrocarbons recovery, chlor-alkali industries, acid gasscrubbing, production of road de-icing salts, and combinations thereof.17. The method of claim 16, wherein step (a) (i) further comprises thestep of mixing said de-oiled water stream with a calcium source; whereinsaid calcium source is selected from the group consisting of dolime,calcium oxide, calcium hydroxide, and combinations thereof.
 18. Themethod of claim 16, further comprises the steps of: (a) feeding at leasta portion of said de-scaled water stream to a Recycle-Brine Multi-StageFlash (RB-MSF) desalination train, wherein said RB-MSF desalinationtrain includes only a heat recovery section, to produce said distillatestream and said de-scaled reject brine; and (b) mixing at least aportion of said de-scaled reject brine with said de-scaled water streamto form a recycle brine stream, and feeding said recycle brine stream tosaid RB-MSF desalination train to produce said distillate stream andsaid de-scaled reject brine; wherein a portion of said de-scaled rejectbrine is discharged from said RB-MSF train at a level not exceeding250,000 mg/L of total dissolved solids (TDS).